Process for production of jet fuels from isomerization and hydrocracking

Abstract

A method separates a liquid hydrocracking stream from a liquid hydroisomerization stream, whereby heat in the hydrocracking stream can be preserved. Preserving heat in the hydrocracking stream avoids the need to reheat the hydrocracking stream prior to product fractionation. In particular, kerosene in the hydrocracking stream is not cooled with the hydroisomerized stream and then reheated in fractionation to distill kerosene range hydrocarbons from the diesel range hydrocarbons.

Description

[0001] Priority declaration

[0002] This application claims priority to U.S. Provisional Patent Application Serial No. 63 / 404,525 filed on September 7, 2022, the entire contents of which are incorporated herein by reference. Technical Field

[0003] The art is the production of hydrocarbons useful as aviation fuel from hydrocarbon feedstocks. In particular, the art may relate to the production of aviation fuel from renewable feedstocks such as triglycerides and free fatty acids found in materials such as plant and animal fats and oils. Background Art

[0004] As the demand for fuel increases worldwide, the production of fuels from sources other than crude oil and the blending of components from sources other than crude oil are also increasingly valued. These sources are commonly referred to as biorenewable sources, including but not limited to vegetable oils, such as corn oil, rapeseed oil, canola oil, soybean oil; microbial oils, such as algae oil; animal fats, such as inedible tallow; fish oils and various waste streams, such as yellow and brown grease and sewage sludge. The common feature of these sources is that they are composed of glycerides and free fatty acids (FFA). Triglycerides and FFAs all contain aliphatic carbon chains with 8 to 24 carbon atoms. The aliphatic carbon chains in triglycerides or FFAs can be fully saturated or mono-, di- or polyunsaturated.

[0005] Hydrotreating can include methods of converting hydrocarbons into more valuable products in the presence of a hydrotreating catalyst and hydrogen. Hydrotreating is a method of contacting hydrogen with hydrocarbons in the presence of a hydrotreating catalyst that is primarily active in removing heteroatoms, such as sulfur, nitrogen, oxygen, and metals, from the hydrocarbon feedstock. In hydrotreating, hydrocarbons with double and triple bonds, such as olefins, can be saturated.

[0006] The production of hydrocarbon products in the diesel boiling range can be achieved by hydrotreating biorenewable feedstocks. Biorenewable feedstocks can be hydrotreated by hydrotreating to deoxygenate, decarboxylate and / or decarbonylate oxygenated hydrocarbons. Decarboxylation and decarbonylation remove carbon from the alkane molecules; deoxygenation does not. Hydrotreating can be followed by hydroisomerization to improve the cold flow properties of the product diesel and jet fuels. Hydroisomerization or hydrodewaxing is a hydrotreating process that increases alkyl branching on the hydrocarbon backbone in the presence of hydrogen and a hydroisomerization catalyst to improve the cold flow characteristics of the hydrocarbons. Hydroisomerization includes hydrodewaxing herein.

[0007] Hydrocracking is a hydroprocessing process in which hydrocarbons are cracked into lower molecular weight hydrocarbons in the presence of hydrogen and a hydrocracking catalyst. Depending on the desired output, a hydrocracking unit may contain one or more catalyst beds that may be identical or different.

[0008] When making jet fuel from triglycerides (also known as "fats"), some degree of hydrocracking and isomerization is required to meet the jet fuel specifications outlined in ASTM D7566 Annex 2 and ASTM D1655. The key jet fuel specifications in D7566 are: a freezing point not higher than -40°C (ASTM D5972, D7153 or D7154), a density not higher than 772 kg / m 3 (ASTM D1298 or D4052), T10 less than 205°C (ASTM D86), and final boiling point (FBP) less than 300°C (ASTM D86). Larger molecules that do not meet these jet fuel specifications are primarily hydrocracking to meet these specifications, which necessarily results in low yields in the manufacturing process and low fuel energy density, which is undesirable. Jet fuel is valued for its high energy per unit volume.

[0009] Carbon intensity is a term referring to the number of moles of carbon dioxide produced to make one mole of fuel. Combustion of hydrocarbons (such as heating a hydrocarbon stream in a hydroprocessing unit) increases carbon intensity. It is desirable to provide renewable fuels from a process that reduces carbon concentration by reducing heating requirements. Summary of the invention

[0010] It has been discovered that by separating a liquid hydroisomerization stream from a liquid hydrocracked stream, heat can be conserved in the hydrocracked stream. Conserving heat in the hydrocracked stream avoids reheating the hydrocracked stream prior to product fractionation. In particular, kerosene in the hydrocracked stream is not cooled with the hydroisomerized stream and then reheated in the fractionation to distill kerosene range hydrocarbons from diesel range hydrocarbons. BRIEF DESCRIPTION OF THE DRAWINGS

[0011] Figure 1 is a schematic process flow diagram of the present disclosure.

[0012] Figure 2 yes Figure 1 Schematic process flow diagram of an alternative embodiment of .

[0013] Figure 3 yes Figure 1 A schematic process flow diagram of another alternative embodiment of .

[0014] Figure 4 yes Figure 1 Schematic process flow diagram of another embodiment of the present invention.

[0015] definition

[0016] The term "communication" means operatively allowing material flow between enumerated components.

[0017] The term "downstream communication" means that at least a portion of the material flowing to the body in the downstream communication can operatively flow from the object in communication therewith.

[0018] The term "upstream communication" means that at least a portion of the material flowing out of the body in the upstream communication can be operatively flowed to an object in communication therewith.

[0019] The term "direct communication" means that a stream from an upstream component enters a downstream component without passing through a fractionation or conversion unit and without undergoing a composition change due to physical fractionation or chemical conversion.

[0020] The term "indirect communication" means that a stream from an upstream component passes through a fractionation or conversion unit and enters a downstream component where it undergoes a composition change due to physical fractionation or chemical conversion.

[0021] The term "bypass" means that the object loses downstream communication with the bypassed body at least within the scope of the bypass.

[0022] The term "tower" means one or more distillation towers for separating one or more components with different volatilities. Unless otherwise specified, each tower includes a condenser on the top of the tower for condensing a portion of the overhead stream and refluxing it back to the top of the tower, and a reboiler at the bottom of the tower for vaporizing a portion of the bottom stream and sending it back to the bottom of the tower. The feed to the tower can be preheated. The top pressure is the pressure of the overhead vapor at the vapor outlet of the tower. The bottom temperature is the liquid bottom outlet temperature. The top line and the bottom line refer to the net line from any reflux or reboil to the tower downstream. The stripping tower can omit the reboiler at the bottom of the tower, but provide heating requirements and separation power for the liquefied inert medium (such as steam). The stripping tower is usually fed from the top tray and the main product is taken out from the bottom.

[0023] As used herein, the term "component-rich stream" refers to a rich stream exiting a vessel having a greater concentration of a component than the feed to the vessel.

[0024] As used herein, the term "component-lean stream" means that the lean stream exiting a vessel has a lesser concentration of a component than the feed to the vessel.

[0025] As used herein, the term "boiling temperature" means the atmospheric equivalent boiling point (AEBP) calculated from the observed boiling temperature and the distillation pressure, as calculated using the formula provided in ASTM D86 or ASTM D2887.

[0026] As used herein, the term "true boiling point" (TBP) means a test method for determining the boiling point of a substance in accordance with ASTM D-2892 for producing standardized masses of liquefied gases, distillate fractions, and residues for which analytical data are available, and for determining the yields of the above fractions by both mass and volume from which a plot of distillation temperature versus mass % is obtained in a column having a reflux ratio of 5:1 using fifteen theoretical plates.

[0027] As used herein, the term "T5" or "T95" means the temperature at which 5 mass percent or 95 mass percent of a sample boils, respectively, using ASTM D-86 or TBP, as the case may be.

[0028] 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.

[0029] As used herein, the term "final boiling point" (FBP) means the temperature at which a sample completely boils using ASTM D2887, ASTM D-86, or TBP, as appropriate.

[0030] As used herein, the term "diesel boiling range" means that using the TBP distillation process, hydrocarbons boil within the range of IBP or T5 and "diesel cut points": IBP is between 125°C (257°F) and 175°C (347°F), T5 is between 150°C (302°F) and 200°C (392°F), and "diesel cut points" include T95 between 343°C (650°F) and 399°C (750°F).

[0031] As used herein, the term "diesel conversion" means the conversion of a feed boiling above the diesel cut point to a material boiling at or below the diesel cut point in the diesel boiling range.

[0032] As used herein, the term "separator" means a vessel having an inlet and at least one overhead vapor outlet and one bottom liquid outlet, and may also have an outlet for an aqueous stream from a boot. A flash tank is a type of separator that may be connected downstream to a separator that may operate at a higher pressure.

[0033] As used herein, the term "predominantly" or "majority" means greater than 50%, suitably greater than 75%, and preferably greater than 90%.

[0034] As used herein, the term “C x " shall be understood to refer to molecules having the number of carbon atoms indicated by the subscript "x". Similarly, the term "C x-" refers to molecules containing less than or equal to x, and preferably x and fewer carbon atoms. The term "C x +" refers to molecules having greater than or equal to x, and preferably x and more carbon atoms.

[0035] As used herein, the term "carbon number" refers to the number of carbon atoms per hydrocarbon molecule, and typically a paraffin molecule. DETAILED DESCRIPTION

[0036] exist Figure 1 In, according to an exemplary embodiment, a method 10 for treating a hydrocarbon feedstock is shown. Preferably, the hydrocarbon feedstock is a bio-renewable hydrocarbon feedstock. A feed line 12 delivers a hydrocarbon stream of a fresh bio-renewable feedstock to a feed buffer tank 14. The bio-renewable feedstock can be mixed with a mineral feed stream, but preferably comprises a major bio-renewable feedstock. Mineral feedstock is a conventional feed derived from crude oil extracted from underground. The bio-renewable feedstock can have a nitrogen concentration of 50wppm to 2000wppm. The bio-renewable feedstock can have a high oxygen content of up to 10% by weight or more. The bio-renewable feedstock can also contain 1wppm to 500wppm sulfur, typically not more than 200wppm sulfur.

[0037] A variety of different biorenewable raw materials can be applicable to method 10. The term "biorenewable raw material" is intended to include raw materials other than those obtained from crude oil. Biorenewable raw materials can include any of those raw materials comprising at least one of glycerides and free fatty acids. Most glycerides are triglycerides, but monoglycerides and diglycerides may also be present and processed. Free fatty acids can be obtained from phospholipids, which can provide phosphorus in the raw materials. Examples of these biorenewable raw materials include, but are not limited to, linseed oil, canola oil, corn oil, soybean oil, rapeseed oil, soybean oil, rapeseed oil, tall oil, sunflower oil, hempseed oil, olive oil, linseed oil, coconut oil, babassu oil, castor oil, peanut oil, palm oil, mustard oil, tallow, yellow and brown grease, lard, whale oil, fat in milk, fish oil, algae oil, sewage sludge, etc. Additional examples of biorenewable feedstocks include non-edible vegetable oils from the group comprising Jatropha curcas (Ratanjot, Wild Castor, Jangli Erandi), Madhucaindica (Mohuwa), Pongamia pinnata (Karanji, Honge), Calophylluminophyllum, Moringa oleifera, and Azadirachtaindica (Neem). Triglycerides and FFAs of typical plant or animal fats contain aliphatic hydrocarbon chains having 8 to 30 carbon atoms in their structure. Biorenewable feedstocks may also include biomass pyrolysis oils and Fischer-Tropsch waxes. As will be appreciated, biorenewable feedstocks may include a mixture of one or more of the foregoing examples. The biorenewable feedstock may be pretreated to remove contaminants and filtered to remove solids.

[0038] The hydrocarbon stream in feed line 12 may flow from feed buffer tank 14 via a feed pump after injecting a sulfiding agent in line 15, and is mixed with a recycled hydroprocessing hydrogen stream in hydroprocessing hydrogen line 20 to provide a combined hydrocarbon stream in line 24. The combined hydrocarbon stream in line 24 is mixed with a hydroprocessing recycle stream in recycle line 16 to provide a hydroprocessing feed hydrocarbon stream in hydroprocessing feed line 26. The recycle to feed ratio may be 1:1 to 5:1. The hydroprocessing feed hydrocarbon stream in line 26 may be preheated in combined feed exchanger 22 by heat exchange with a hydroprocessing stream in hydroprocessing line 32 and possibly then in fired heater 23. The heated hydroprocessing feed hydrocarbon stream in hydroprocessing feed line 26 may then be charged to hydroprocessing reactor 28.

[0039] The hydroprocessing reactor 28 may include a guard bed reactor or guard bed 27. Figure 1In the embodiment of the present invention, the hydroprocessing reactor 28 includes a guard bed 27. The guard bed reaction temperature may range between 246°C (475°F) and 343°C (650°F), and suitably between 288°C (550°F) and 304°C (580°F). The reaction temperature is low enough to prevent olefin polymerization in the FFA, but high enough to promote olefin saturation, hydrodemetallization, hydrodeoxygenation, hydrodesulfurization, and hydrodenitrogenation reactions. The hydrodeoxygenation reaction preferably minimizes the hydrodecarbonylation reaction and the hydrodecarboxylation reaction to retain the carbon atoms on the paraffin chain.

[0040] The guard bed 27 may comprise a base metal catalyst on a support. Base metals that may be used in the 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, wherein the metal in the metal sulfide is selected from one or more of the listed base metals. The hydroprocessing feed 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 another embodiment, the guard bed catalyst may comprise a second metal, wherein the second metal comprises one or more of the following metals: tin, indium, ruthenium, rhodium, rhenium, osmium, iridium, germanium, lead, cobalt, gallium, zinc, and thallium. A nickel-molybdenum catalyst on alumina may be a suitable catalyst in the guard bed 27. Although in Figure 1 Only one guard bed is shown in FIG. 2 , but multiple guard beds may be included in the hydroprocessing reactor 28 , such as 2, 3 or more, and a hydrogen quench from the hydrogen manifold 18 may be injected at inter-bed locations to control the temperature exotherm.

[0041] The contacted hydrocarbon stream is discharged from guard bed 27. In guard bed 27, most of the hydrodemetallization and hydrodeoxygenation reactions will occur along with some hydrodenitrogenation and hydrodesulfurization. Metals removed from the biorenewable feedstock will include alkali and alkaline earth metals and phosphorus. If the guard bed has a dedicated reactor vessel, the contacted hydrocarbon stream will be discharged from the guard bed reactor. However, in Figure 1 In the embodiment, guard bed 27 is contained in hydroprocessing reactor 28, so the contacted stream will receive hydrogen quench from hydrogen manifold 18 and enter hydroprocessing catalyst bed 29.

[0042] In the hydrotreating reactor 28, the contacted hydrocarbon stream is contacted with a hydrotreating catalyst in a hydrotreating catalyst bed 29 under hydrotreating conditions in the presence of hydrogen to saturate the olefinic or unsaturated portions of the n-paraffin chains in the biorenewable feedstock. The hydrotreating catalyst also catalyzes hydrodeoxygenation reactions, including hydrodecarboxylation and hydrodecarbonylation reactions, to remove oxygenate functional groups from hydrocarbon molecules in the biorenewable feedstock, which are converted into water and carbon oxides. The hydrotreating catalyst also catalyzes the hydrodesulfurization of organic sulfur and the hydrodenitrogenation of organic nitrogen in the biorenewable feedstock. Basically, the hydrotreating reaction removes heteroatoms from the hydrocarbons and saturates the olefins in the feed stream.

[0043] The hydroprocessing catalyst may be provided in one, two or more beds and employ an inter-bed hydrogen quench stream from the hydrogen quench stream. A recycle hydrogen quench stream taken from recycle hydrogen line 19 in hydrogen manifold line 18 may be provided to the hydroprocessing reactor 28 for inter-bed quenching. Figure 1 Three hydrotreating catalyst beds 29 are shown, but one or more are contemplated.

[0044] The hydroprocessing 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 precious metals dispersed on a high surface area support. Non-limiting examples of precious metals include platinum and / or palladium dispersed on an alumina support such as γ-alumina. Suitable hydroprocessing catalysts include BDO 200 or BDO 300 or BDO 400 purchased from UOP LLC, Des Plaines, Illinois, USA. The hydroprocessing reaction temperature may range between 271°C (520°F) and 427°C (800°F), and preferably between 304°C (580°F) and 400°C (752°F). Typically, hydroprocessing conditions include a pressure of 700 kPa (100 psig) to 21 MPa (3000 psig).

[0045] A hydrotreated stream comprising a hydrocarbon fraction having a significant normal paraffin concentration is produced in hydrotreating line 32 from hydrotreating reactor 28. The oxygenate concentration in the hydrocarbon fraction is substantially zero, while the olefin concentration is significantly reduced relative to the contacted biorenewable feed stream. The organic sulfur concentration in the hydrocarbon fraction does not exceed 500 wppm, and the organic nitrogen concentration in the hydrocarbon fraction is less than 10 wppm.

[0046] The hydrotreated stream in hydrotreating line 32 may first flow to combined isomerization feed exchanger 34 to heat the hydroisomerization feed stream in hydroisomerization feed line 44 and cool the hydrotreated stream. As previously described, the cooled hydrotreated stream in hydrotreating line 32 may then be heat exchanged with the combined hydrocarbon stream in line 24 in combined feed heat exchanger 22 to further cool the cooled hydrotreated stream in hydrotreating line 32 and heat the hydrotreated feed hydrocarbon stream in hydrotreating line 32. The twice cooled hydrotreating steam in hydrotreating line 32 may then be further cooled in combined feed exchanger 22 by heat exchange with the combined hydrocarbon stream in combined feed line 24 to heat the combined hydrocarbon stream and cool the hydrotreating stream in hydrotreating line 32. The twice cooled hydrotreating stream may be even further cooled before separation, perhaps to produce steam.

[0047] The cooled hydrocarbon stream may be separated in a hydrotreating separator 36, which may include an enhanced hot separator (EHS), with the aid of a stripping gas fed in stripping line 39 from isomerization column overhead line 58. The hydrotreated stream is separated to provide a hydrotreated vapor stream in hydrotreating column overhead line 38 and a hydrotreated liquid stream in hydrotreating column bottoms line 40 having a lower oxygen concentration than the hydrotreated feed hydrocarbon stream in line 26. The hydrotreating separator 36 may be a high pressure stripping column. In the hydrotreating separator 36, the hydrotreating stream from the hydrotreating line 32 flows downwardly through the column where the hydrotreating stream is partially stripped of hydrogen, carbon dioxide, carbon monoxide, water vapor, propane, hydrogen sulfide, and phosphine, which are potential isomerization catalyst poisons, by contact with a stripping gas from stripping line 39. The stripping gas may include make-up hydrogen that has passed through isomerization reactor 48 and isomerization separator 56, as described below.

[0048] The stripping gas in stripping line 39 enters hydrotreating separator 36 below the inlet of the hydrotreated stream in hydrotreating line 32. Hydrotreating separator 36 may include internal components such as trays or packing located between the inlet of the hydrotreated stream in line 32 and the inlet of the stripping gas in stripping line 39 to facilitate stripping of the hydrotreated stream. The stripping gas, including the stripped gas, exits as a hydrotreated vapor stream in hydrotreating overhead line 38 extending from the top of hydrotreating separator 36, mixes with a cold aqueous stream in cold aqueous line 63 from a storage tank of cold separator 62, is cooled, and mixes with an isomerized liquid stream in isomerization bottoms line 60, and enters cold separator 62.

[0049] The hydrotreating separator 36 operates at 177°C (350°F) to 371°C (700°F), and preferably operates at 232°C (450°F) to 315°C (600°F). The hydrotreating separator 36 may operate at a slightly lower pressure than the hydrotreating reactor 28 to account for the pressure drop through the intervening equipment. The hydrotreating separator 36 may operate at a pressure between 3.4 MPa (gauge) (493 psig) and 20.4 MPa (gauge) (2959 psig). The temperature of the hydrotreating vapor stream in the hydrotreating separator overhead line 38 may be the operating temperature of the hydrotreating separator 36.

[0050] The hydrotreated liquid stream, which may have been stripped, is collected at the bottom of the hydrotreating separator 36 and flows in the hydrotreating bottoms line 40. The liquid hydrotreating stream contains diesel range materials and has a high paraffin concentration if the hydrocarbon feed contains a biorenewable feedstock. The liquid hydrotreating stream in the hydrotreating separator bottoms line 40 can be divided into two streams: a hydroisomerization feed stream withdrawn in the hydroisomerization feed line 42 and a recycle hydrotreating stream withdrawn in the recycle line 16, both of which are taken from the liquid hydrotreating stream in the hydrotreating bottoms line 40. The recycle hydrotreating stream in the recycle line 16 can be pumped as previously described and combined with the combined hydrocarbon stream in the line 24.

[0051] Although the liquid hydroprocessed stream may provide a desired product (such as a transportation fuel) in the hydroprocessing column bottoms line 40 due to its high concentration of normal paraffins, it will have poor cold flow properties and high FBP, making it fail to meet jet fuel specifications. Therefore, in order to improve the cold flow properties and reduce the FBP, the hydroprocessed liquid stream may be hydroisomerized.

[0052] The make-up hydrogen in make-up line 41 may be compressed in make-up gas compressor 45 to provide compressed make-up gas in compressed make-up gas header 47. A hydroisomerization make-up gas stream is withdrawn from make-up gas header 47 in line 43 and mixed with a hydroisomerization feed stream in line 42 to provide a combined hydroisomerization feed stream in combined hydroisomerization feed line 44. The combined hydroisomerization feed stream in combined hydroisomerization feed line 44 may be heated in hydroisomerization feed exchanger 34 by heat exchange with a hydrotreating stream in hydrotreating line 32. The combined hydroisomerization feed stream may be heated in hydroisomerization feed heater 46 to bring the combined hydroisomerization feed stream to a hydroisomerization temperature before the combined hydroisomerization feed stream is charged to hydroisomerization reactor 48.

[0053] The hydroisomerization (including hydrodewaxing) of normal hydrocarbons in the hydroisomerization reactor 48 can be accomplished by one or more hydroisomerization catalyst beds, and the hydroisomerization can be operated in a cocurrent operation mode. Fixed bed, trickle bed downflow or fixed bed submerged upflow modes are all suitable.

[0054] The hydroisomerization catalyst comprises a dehydrogenation metal, a molecular sieve and a metal oxide binder. The hydroisomerization catalyst may comprise a dehydrogenation metal containing a Group VIII metal. The dehydrogenation metal may be selected from platinum, palladium, nickel, nickel-molybdenum sulfide or nickel-tungsten sulfide. Preferably, the dehydrogenation metal is selected from platinum or nickel-tungsten sulfide. The concentration of the dehydrogenation metal on the hydroisomerization catalyst may be 0.05 wt % to 5 wt % based on the transition metal composition.

[0055] The dehydrogenation metal distributed between the molecular sieve and the binder is 40% to 65% by weight (preferably 45% to 60% by weight) of the metal distributed on the molecular sieve, and 40% to 65% by weight (preferably 45% to 60% by weight) of the metal distributed on the binder. The related benefits of the hydroisomerization catalyst are high activity and selectivity for hydroisomerization. In other embodiments, the hydroisomerization catalyst also contains less than 0.5% by weight of carbon, which has the related benefits of high activity and selectivity for hydroisomerization.

[0056] In one embodiment, the hydroisomerization 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, such as 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, ferrierite, mordenite, ZSM-5 or zeolite beta, with the associated benefit of being active in the hydroisomerization of linear hydrocarbons.

[0057] The metal oxide binder may be selected from alumina, silica, silica-alumina and titania or mixtures thereof.Preferably, the metal oxide binder is alumina, and preferably it is gamma-alumina.

[0058] The hydroisomerization catalyst may comprise a molecular sieve having an AEL topology, and more specifically SAPO-11. Most of the acid sites on SAPO-11 are weak to moderate acid sites. More specifically, at least 50% of the total acidity on SAPO-11 is weak acidity, and at least 60% to 80% of the external acidity on SAPO-11 is weak acidity.

[0059] Hydroisomerization catalysts typically comprise particles having a diameter of 1 mm to 5 mm. The catalyst preparation typically involves forming a stable porous support followed by impregnation with an active metal. The stable porous support typically comprises a metal oxide and a molecular sieve, which may be a zeolite. A stable support having a high porosity is prepared to ensure 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, may be a suitable hydroisomerization catalyst.

[0060] The hydroisomerization conditions typically include a temperature of 150°C (302°F) to 450°C (842°F) and a pressure of 1724 kPa (absolute) (250 psia) to 13.8 MPa (absolute) (2000 psia). In another embodiment, the hydroisomerization conditions include 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 0.5 hr -1 Up to 3 hours -1 LHSV, and 337Nm 3 / m 3 (2,000scf / bbl) to 2,527Nm 3 / m 3 Hydrogen rate to oil (15,000 scf / bbl).

[0061] The hydroisomerization stream from the hydroisomerization reactor 48 in the hydroisomerization line 50 is a stream rich in branched paraffins. Preferably, the hydroisomerization stream is mainly a branched paraffin stream. It is envisioned that the hydroisomerization effluent may contain a total alkane content of 80 mass%, 90 mass% or 95 mass% branched paraffins. The hydroisomerization conditions in the hydroisomerization reactor 48 are selected to avoid undesirable cracking, so the main product in the hydroisomerization stream in the hydroisomerization line 50 is branched paraffins. By avoiding undesirable cracking, the hydroisomerization stream in the hydroisomerization line 50 has a similar composition in terms of carbon number that is close to and only slightly less than the hydroisomerization feed stream in the hydroisomerization feed line 42. The optimal amount of remaining normal paraffins in the pipeline 50 depends on the selectivity of the hydroisomerization catalyst, but can generally be between 1 weight % and 7 weight %.

[0062] The hydroisomerized stream from the hydroisomerization reactor 48 in the hydroisomerization line 50 flows to the hydroisomerize exchanger 52 to exchange heat with the stripper liquid hydroisomerized stream in the cold bottoms line 70 to cool the hydroisomerized stream before it is mixed with the vapor hydrocracked stream in the line 182 to provide the mixed hydroisomerized stream in the line 54. The mixed hydroisomerized stream in the line 54 is further cooled in the hydroisomerization cooler 55 and fed to the hydroisomerization separator 56 to be separated into a liquid hydroisomerized stream and a vapor hydroisomerized stream. Internal packing may be located at the top of the hydroisomerization separator 56 to ensure that the liquid component is blocked from leaving the hydroisomerization tower top line 58. The vaporous hydroisomerize stream in hydroisomerization overhead line 58 extending from the top of the hydroisomerization separator 56 may be compressed in compressor 59 to provide stripping gas in stripping line 39 to the hydrotreating separator 36 .

[0063] The liquid hydroisomerized stream in the hydroisomerization column bottoms line 60 extending from the bottom of the hydroisomerization separator 56 may be fed to the stripper 100 or to the product fractionator 120 separated from the hydrocracked liquid stream to be described below. In one embodiment, the liquid hydroisomerized stream in the hydroisomerization column bottoms line 60 may be pumped to a cold separator 62 for further separation along with the cooled vapor hydrotreated stream in line 38 and the cold aqueous stream in the cold aqueous line 63 pumped from around the storage tank of the cold separator 62. The cold aqueous stream in line 63 may be combined with the vapor hydrotreated stream in line 38 to provide a relatively cold hydrotreated stream in line 61. The cooler hydrotreated stream in line 61 may be cooled in cooler 64 and combined with the liquid hydroisomerize stream in line 60 to provide a cold separator feed stream in line 66 and fed to cold separator 62. A cold aqueous stream in cold aqueous line 63 is added to cold separator feed line 66 via the liquid hydroisomerize stream in line 60 to dissolve salts that may be present in the liquid hydrocarbons in cold separator 62.

[0064] In cold separator 62, vapor components of the hydroisomerize liquid stream and the hydrotreating vapor stream will separate and rise to provide a cold vapor hydroisomerize stream in cold overhead line 68 and a stripper liquid hydroisomerize stream in cold bottoms line 70 and a cold aqueous stream taken from storage tank in cold aqueous line 63. The cold vapor hydroisomerize steam in cold overhead line 68 may be split between a recycle hydrogen stream in line 19 and a net cold vapor hydroisomerize stream in net cold overhead line 72. The recycle hydrogen stream in line 19 is compressed in a recycle gas compressor and recycled to the hydrotreating reactor 28 in manifold line 18 for inter-bed quenching and recycled in hydrotreating hydrogen line 20 to the hydrocarbon stream in feed line 12.

[0065] The net cold vapor hydroisomerize stream in net cold overhead line 72 may be passed through a trayed or packed purge scrubber 74 wherein the net cold vapor hydroisomerize stream is scrubbed with a scrubbing liquid, such as an aqueous solution, fed through a scrubbing liquid line 76 to absorb the acid gases containing hydrogen sulfide and carbon dioxide by extracting the acid gases into the aqueous solution. Preferred scrubbing liquids include Selexol TM , which is available from UOP LLC of Des Plaines, Illinois, and amines such as alkanolamines, including diethanolamine (DEA), monoethanolamine (MEA), methyldiethanolamine (MDEA), diisopropanolamine (DIPA) and diglycolamine (DGA). Other scrubbing liquids may be used instead of or in addition to the preferred amines. The lean scrubbing liquid contacts the cold vapor stream and absorbs the acidic gas contaminants. The resulting "mildened" cold vapor stream is taken from the top outlet of the purge scrubber 74 in the purge scrubber top line 78, and the rich acid gas scrubbing liquid is taken from the bottom of the purge scrubber 74 at the bottom outlet of the purge scrubber 74 in the purge scrubber bottom line 80. The used scrubbing liquid from the bottom of the tower can be regenerated in the scrubbing liquid line 76 and recycled back to the purge scrubber 74. A scrubbed hydrogen-rich stream exits the purge scrubber 74 in purge scrubber overhead line 78 and may be sent to a pressure swing adsorption unit 82 or other hydrogen recovery device to produce high purity hydrogen in line 84 .

[0066] The purge scrubber 74 may be operated at a gas inlet temperature between 38°C (100°F) and 66°C (150°F) and an overhead pressure of 3 MPa (gauge) (435 psig) to 20 MPa (gauge) (2900 psig). Suitably, the purge scrubber 74 may be operated at a temperature of 40°C (104°F) to 125°C (257°F) and a pressure of 1200 kPa to 1600 kPa. The temperature of the net cold vapor hydroisomerization stream 72 to the purge scrubber 74 may be between 20°C (68°F) and 80°C (176°F), and the temperature of the scrubber liquid stream in the scrubber liquid line 76 may be between 20°C (68°F) and 70°C (158°F).

[0067] The liquid hydroisomerized fuel component in the stripper liquid hydroisomerization stream from the liquid hydroisomerization stream and the steam hydrotreated stream will leave the cold separator in the cold hydroisomerization bottoms line 70. The stripper liquid hydroisomerization stream in the cold hydroisomerization bottoms line 70 contains diesel and jet boiling range fuels and other hydrocarbons such as propane and naphtha. The cold aqueous stream can be collected from the storage tank of the cold separator in the cold aqueous line 63.

[0068] In an embodiment, the stripper liquid hydroisomerize stream in the cold bottoms line 70 may be stripped to remove hydrogen sulfide and other gases in a cold stripper 86. The stripper liquid hydroisomerize stream in the cold bottoms line 70 may be heated by heat exchange with the hydroisomerize stream in the hydroisomerize line 50 in the hydroisomerize exchanger 52 to heat the cold liquid hydroisomerize stream and feed to the cold stripper 86.

[0069] In an embodiment, the cold stripper 86 can be part of a dual stripping vessel 100, which includes a cold stripper 86 and a hot stripper 102 separated from each other by an intermediate wall 103. A stripping medium such as steam, which is an inert gas, from a stripping medium line 89 can be used to strip light gases from the stripper liquid hydroisomerization stream in line 70. The cold stripper 86 provides an overhead stripping stream of naphtha, LPG, hydrogen, hydrogen sulfide, steam and other gases in a stripper overhead line 87, and a fractionator hydroisomerization stream in a cold stripper bottom line 90. The overhead stripping stream in the overhead line 87 can be condensed and separated by cooling in a stripping receiver 95. The stripper overhead line 88 from the receiver 95 can carry the stripper overhead stream to the exhaust scrubber 140. The unstabilized liquid naphtha from the bottom of receiver 95 can be split to provide a reflux stream to cold stripper 86 and a stripper liquid overhead stream that can be sent to debutanizer 170 for naphtha and LPG recovery in stripper receiver bottoms line 96. A sour water stream can be collected from a storage tank of overhead receiver 95.

[0070] The cold stripper 86 can be operated at a top pressure of 0.35 MPa (gauge) (50 psig), preferably not less than 0.70 MPa (gauge) (100 psig) to not more than 2.0 MPa (gauge) (290 psig). The temperature in the overhead receiver 95 is in the range of 38° C. (100° F.) to 66° C. (150° F.), and the pressure is substantially the same as the top pressure of the cold stripper 86. It is contemplated that the cold stripper 86 and the hot stripper 102 can be two completely separate vessels.

[0071] The fractionator liquid hydroisomerization stream in the cold stripper bottoms line 90 may be heated and fed to the product fractionator 120. The fractionator liquid hydrocracked stream in the hot stripper bottoms line 106 may also be heated and fed to the product fractionator 120. The fractionator liquid hydrocracked stream in the hot stripper bottoms line 106 is a liquid stream separated from the fractionator liquid hydroisomerization stream (which is also a liquid stream) in the cold stripper bottoms line 90. The fractionator liquid hydrocracked stream in the hot stripper bottoms line 106 may be fed to the product fractionator 120 separately or together with the fractionator liquid hydroisomerization stream in the cold stripper bottoms line 90. However, the fractionator liquid hydrocracked stream in the hot stripper bottoms line 106 and the fractionator liquid hydroisomerization stream in the cold stripper bottoms line 90 are generated separately. These streams are removed and kept separate until stripping or fractionation is used to conserve heat in the fractionator liquid hydrocracked stream so that less heat is required to boil the kerosene range boiling components, thereby separating them from the diesel components in the product fractionator 120.

[0072] The diesel stream in the bottom line 124 is taken from the bottom of the product fractionator 120. The hydrocracking feed stream in the line 126 is taken from the diesel stream in the bottom line 124 from the product fractionator 120. The product fractionator 120 can be reboiled by heat exchange with a suitable hot stream or in the fired heater 121 to provide the heat required for distillation. Alternatively, the tower can be heated using a stripping medium such as steam as an inert gas from the stripping medium line. The reboiled stream is taken to the fired heater 121 and boiled back to the product fractionator 120. The diesel product stream can be taken to the diesel pool in the diesel product line 125 and can be green diesel. The diesel stream in the distillation tower bottom line 124 can be a diesel stream with T5 of 230°C (446°F) to 296°C (590°F) and T90 of 343°C (650°F) to 399°C (750°F).

[0073] The product fractionation column 120 provides an overhead gaseous stream of naphtha in an overhead line 122. The fractionation overhead stream can be fully condensed and separated from water in a fractionation receiver 130. Unstabilized liquid naphtha from the bottom of receiver 130 in a fractionator overhead liquid line 132 can be combined with the naphtha stream in line 176 while a condensed reflux stream is refluxed to the column. A sour water stream can be collected from a storage tank of the distillation receiver 130.

[0074] A kerosene stream may be withdrawn from the side of the product fractionation column 120 in a side line 134. The kerosene stream withdrawn in the side line 134 may be stripped in a kerosene stripper 136 to drive off lower boiling point materials which are returned to the product fractionation column 120 at a higher elevation in an overhead kerosene line 135. A stripped bottoms kerosene stream is produced in a bottoms kerosene line 137 from which a boiling stream is reboiled and fed back to the kerosene stripper 136 and a jet fuel product stream is withdrawn in line 138. The jet fuel product stream in line 138 meets the jet fuel specifications according to ASTM D86 and may be a green jet fuel stream withdrawn from the bottom of the kerosene stripper 136. The jet fuel product stream in line 138 may be cooled and sent to a jet fuel pool.

[0075] Optionally, a light diesel stream can be withdrawn in a second side line and stripped in a side diesel stripper not shown. In addition, heat can be removed from the product fractionator in a pump-cycle steam generator 139.

[0076] The product fractionator 120 may be operated at a bottom temperature between 149°C (300°F) and 288°C (550°F), preferably not more than 260°C (500°F), and an overhead pressure of 0.35 MPa (gauge) (50 psig), preferably not less than 0.70 MPa (gauge) (100 psig) to not more than 2.0 MPa (gauge) (290 psig). The temperature in the overhead receiver 130 is in the range of 38°C (100°F) to 66°C (150°F), and the pressure is substantially the same as the overhead pressure of the product fractionator 120. It is also contemplated that the product fractionator 120 may provide only a net overhead stream comprising jet fuel in the fractionator overhead liquid line 132, with naphtha and lighter streams being taken in the fractionator receiver overhead line 131. In such a case, the compressor 133 would not be required.

[0077] The overhead stripping stream of naphtha, LPG, hydrogen, hydrogen sulfide, steam and other gases in the stripper overhead line 88 may be passed through a trayed or packed off-gas scrubber 140 where the overhead stripping stream is scrubbed with a scrubbing liquid (such as an aqueous solution) fed through a scrubbing liquid line 142 to remove acid gases containing hydrogen sulfide and carbon dioxide by extracting the acid gases into the aqueous solution. Preferred scrubbing liquids include Selexol TM, which is purchased from UOP LLC of Des Plaines, Illinois, and amines such as alkanolamines, including diethanolamine (DEA), monoethanolamine (MEA), methyldiethanolamine (MDEA), diisopropanolamine (DIPA) and diglycolamine (DGA). Other scrubbing liquids may be used instead of or in addition to the preferred amines. The lean scrubbing liquid contacts the overhead stripping stream and absorbs acidic gas pollutants such as hydrogen sulfide and carbon dioxide. The resulting "mildened" overhead stripping stream is taken out from the overhead outlet of the exhaust gas scrubber 140 in the recycle scrubber overhead line 144, and the rich acid gas scrubbing liquid is taken out from the bottom of the tower at the bottom outlet of the recycle scrubber tower 140 in the recycle scrubber bottom line 146. The used scrubbing liquid from the bottom of the tower can be regenerated in the scrubbing liquid line 142 and recycled back to the exhaust gas scrubber 140. A scrubbed hydrocarbon-rich stream is discharged from the off-gas scrubber tower 140 via the off-gas scrubber overhead line 144 and may be sent to a sponge absorber tower 160 for hydrocarbon recovery.

[0078] The off-gas scrubber 140 may be operated at a gas inlet temperature between 38°C (100°F) and 66°C (150°F) and an overhead pressure of 3 MPa (gauge) (435 psig) to 20 MPa (gauge) (2900 psig). Suitably, the off-gas scrubber 140 may be operated at a temperature of 40°C (104°F) to 125°C (257°F) and a pressure of 1200 kPa to 1600 kPa. The temperature of the overhead stripping stream 88 to the off-gas scrubber 140 may be between 20°C (68°F) and 80°C (176°F), and the temperature of the scrubber liquid stream in the scrubber liquid line 142 may be between 20°C (68°F) and 70°C (158°F).

[0079] The sponge absorber tower 160 can receive the washed hydrocarbon-rich stream in the exhaust scrubber overhead line 144. The lean absorbent stream in the lean absorbent line 162 can be fed into the sponge absorber tower 160 through the absorbent inlet. The lean absorbent can include a naphtha stream in the lean absorbent line 162, possibly from the debutanizer bottoms stream in line 176. In the sponge absorber tower 160, the lean absorbent stream and the washed hydrocarbon-rich stream are countercurrently contacted. The sponge absorbent absorbs the LPG hydrocarbons from the clean stripper gaseous stream into the rich absorbent stream.

[0080] The hydrocarbons absorbed by the sponge absorbent include some methane and ethane as well as most of the LPG, C3 and C4 hydrocarbons in the net stripper gas stream and any C5 and C 6+Light naphtha hydrocarbons. The sponge absorber tower 160 operates at a temperature of 34°C (93°F) to 60°C (140°F) and at a pressure substantially the same as or lower than the exhaust gas scrubber 140, thereby reducing friction losses. The sponge absorption exhaust gas stream depleted of LPG hydrocarbons is discharged from the top of the sponge absorption tower 160 through the sponge absorption tower overhead pipeline 164 at the tower top outlet. The sponge absorption exhaust gas stream in the sponge absorber tower overhead pipeline 164 can be transported to a fuel gas header not shown for providing fuel gas requirements. The rich absorbent stream rich in LPG hydrocarbons is discharged from the bottom of the sponge absorber tower 160 at the tower bottom outlet in the rich absorber bottom pipeline 166, which can be fed to the debutanizer 170 via the stripper tower overhead liquid stream in the stripper receiver tower bottom pipeline 96.

[0081] In one embodiment, the debutanizer 170 can fractionate the stripper liquid overhead stream and the rich absorbent stream in the stripper receiver bottoms line 134 into a stream comprising primarily C 5+ The butanizer bottoms stream of hydrocarbons and the butanizer overhead stream comprising LPG hydrocarbons. The butanizer overhead stream in the butanizer overhead line 172 can be completely condensed by reflux to the butanizer 170 and recovering the LPG in the butanizer overhead liquid stream in the butanizer net receiver bottoms line 174. The butanizer overhead liquid stream in the net receiver bottoms line 174 can be taken out as an LPG product stream. The butanizer bottoms stream can be discharged from the bottom of the butanizer 170 in the butanizer bottoms line 176. The reboiled stream taken out from the butanizer bottoms stream in the butanizer bottoms line 175 from the bottom of the butanizer 170 can be boiled in the reboiler line and sent back to the butanizer 170 to provide heat to the tower. Alternatively, a hot inert medium stream such as steam can be fed to the tower 170 to provide heat. The net debutanizer bottoms stream in line 176 comprising naphtha may be split between a lean absorbent stream in lean absorbent line 162 and a product naphtha stream which is cooled and sent to a gasoline pool in line 178 .

[0082] The fractionator bottoms stream in fractionator bottoms line 124 may include diesel boiling range hydrocarbons. Figure 1 In an embodiment, the jet fuel stream in line 138 and the diesel stream in line 125 can be withdrawn at once without recycle. The split point in the product fractionator 120 between the diesel stream in bottoms line 124 and the jet fuel stream in side line 134 can be adjusted to ensure that the jet fuel stream has an appropriate composition to meet the jet fuel specifications, at least after blending, and in particular, meet the jet fuel density specifications. However, because the larger paraffins are concentrated in the fractionator bottoms stream, it is very suitable for hydrocracking into kerosene range hydrocarbons.

[0083] The hydrocracking feed stream in hydrocracking feed line 126 may be stored in a hydrocracking feed buffer tank 127 and pumped to a hydrocracking reactor 150. The hydrocracking reactor 150 is downstream of the hydroisomerization reactor 48. The hydrocracking feed stream may be mixed with a hydrocracking hydrogen stream in line 152 (which is taken from a compressed make-up hydrogen stream in a compressed make-up gas header 47) to provide a combined hydrocracking feed stream in a combined hydrocracking feed line 154, heated by heat exchange with a hydrocracking stream in line 158 in a hydrocracking feed effluent feed heater 155, and charged to the hydrocracking reactor 150 in line 154.

[0084] Hydrocracking reactor 150 can be a fixed bed reactor, which includes one or more containers, a single or multiple catalyst beds in each container, and various combinations of hydrocracking catalysts in one or more containers. Hydrocracking reactor 150 can be operated in a conventional continuous gas phase, moving bed or fluidized bed hydroprocessing reactor.

[0085] The combined hydrocracked streams are hydrocracked in hydrocracked reactor 150 over a hydrocracked catalyst in the presence of a hydrocracked hydrogen stream from hydrocracked hydrogen line 152 to provide a hydrocracked stream. A portion of the hydrocracked feed stream in line 126 may be used as an inter-bed quench to cool the hydrocracked effluent between catalyst beds. In an alternative aspect, recycle hydrogen from recycle hydrogen line 19 may be added between hydrocracked catalyst beds.

[0086] The hydrocracking reactor can provide a total conversion of at least 20% by volume and typically greater than 60% by volume of the hydrocracking feed stream in the hydrocracking feed line 126 to products having a boiling point below the heavy diesel range of 293° C. (560° F.) to 310° C. (590° F.). The hydrocracking reactor 150 can be operated at a partial conversion of more than 30% by volume of the feed or a full conversion of at least 90% by volume based on the total conversion. The hydrocracking reactor 150 can be operated under mild hydrocracking conditions that will provide a total conversion of 20% to 60% by volume, preferably 20% to 50% by volume, of the hydrocracking feed stream to products having a boiling point below the heavy diesel boiling point range.

[0087] The hydrocracking catalyst may use an amorphous silica-alumina base or a zeolite base in combination with one or more Group VIII or Group VIB metal hydrogenation components to selectively produce a balance of light diesel and jet fuel distillates. In another aspect, a catalyst generally comprising any crystalline zeolite cracking base having a Group VIII metal hydrogenation component deposited thereon may be suitable. Additional hydrogenation components may be selected from the VIB group to be combined with the zeolite base. In addition, the hydroisomerization catalyst of the hydroisomerization reactor 48 may be used as a hydrocracking catalyst in the hydrocracking reactor 150, but operates at the high end of the hydroisomerization temperature range.

[0088] Zeolite cracking base is sometimes referred to as molecular sieve in the art, and is generally composed of silica, alumina and one or more exchangeable cations (such as sodium, magnesium, calcium, rare earth metals). It is also characterized by having a relatively uniform crystal pore with a diameter between 4 angstroms and 14 angstroms. Zeolites with a relatively high silica / alumina molar ratio (between 3 and 12) are preferably used. Suitable zeolites found in nature include, for example, mordenite, stilbite, heulandite, ferrierite, ring crystal zeolite, chabazite, erionite and faujasite. Suitable synthetic zeolites include, for example, B, X, Y and L crystal types, such as synthetic faujasite and mordenite. Preferred zeolites are those zeolites with a crystal pore size between 8 angstroms and 12 angstroms, wherein the molar ratio of silica / alumina is 4 to 6. An example of a zeolite falling into the preferred group is a synthetic Y molecular sieve.

[0089] Naturally occurring zeolites are usually present 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 cracking base materials, it is preferred that most or all of the original zeolite monovalent metals are ion exchanged with polyvalent metals and / or with ammonium salts, and then heated to decompose the ammonium ions associated with the zeolite, thereby leaving hydrogen ions and / or exchange sites that are actually decationized by further removing water in their place. Hydrogen or "decationized" Y zeolites of this nature are more specifically described in US 3,100,006.

[0090] Mixed multivalent metal-hydrogen zeolites can be prepared by exchanging ions with ammonium salts, then partially back-exchanging with multivalent metal salts, and then calcining. In some cases, such as in the case of synthetic mordenite, the hydrogen form can be prepared by directly acid treating alkali metal zeolites. In one aspect, preferred cracking bases are those that lack at least 10% by weight and preferably at least 20% by weight of metal cations based on the initial ion exchange capacity. In another aspect, an ideal and stable class of zeolites is one in which the hydrogen ions satisfy at least 20% by weight of the ion exchange capacity.

[0091] The active metal used as the hydrogenation component in the preferred hydrocracking catalyst of the present disclosure is an active metal of Group VIII, i.e., iron, cobalt, nickel, ruthenium, rhodium, palladium, osmium, iridium and platinum. In addition to these metals, other promoters may be used in combination, including Group VIB metals, such as molybdenum and tungsten. The amount of hydrogenation metal in the catalyst may vary over a wide range. In general, any amount between 0.05 wt % and 30 wt % may be used. With respect to noble metals, it is generally preferred to use 0.05 wt % to 2 wt % of noble metals. Noble metals may be preferred as hydrogenation metals on hydrocracking catalysts to provide selectivity to jet fuels, due to the absence of hydrogen sulfide and ammonia, which may deactivate the noble metal catalyst but have been removed upstream of the process.

[0092] The method of incorporating the hydrogenation metal is to contact the base material with an aqueous solution of a suitable compound of the desired metal, wherein the metal is present in cationic form. After addition of the selected hydrogenation metal or metals, the resulting catalyst powder is then filtered, dried, pelletized with added lubricants, binders, etc. as required, and calcined in air at a temperature of, for example, 371° C. (700° F.) to 648° C. (200° F.) to activate the catalyst and decompose the ammonium ions. Alternatively, the base component can be pelletized, and then the hydrogenation component added and activated by calcination.

[0093] The above catalysts can be used in undiluted form, or the powdered catalysts can be mixed with other relatively less active catalysts, diluents or binders such as alumina, silica gel, silica-alumina co-gel, activated clay, etc. in a ratio ranging between 5% and 90% by weight and co-pelletized. These diluents can be used as is, or they can contain a smaller proportion of added hydrogenation metals, such as VIB group and / or VIII group metals. Additional metal-promoted hydrocracking catalysts can also be used in the process of the present disclosure, including, for example, aluminum phosphate molecular sieves, crystalline chromosilicates and other crystalline silicates. Crystalline chromosilicates are more fully described in US 4,363,178.

[0094] By one approach, the hydrocracking conditions may include 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) (400 psig) to 20.7 MPa (gauge) (3000 psig), 0.4 hr -1 To less than 2.5hr -1 Liquid Hourly Space Velocity (LHSV), and 337Nm 3 / m 3 (2,000scf / bbl) to 2,527Nm3 / m 3 Hydrogen rate for oil (15,000 scf / bbl).

[0095] The hydrocracked stream may exit the hydrocracking reactor 150 in hydrocracked line 158. In an embodiment, the hydrocracked stream may be cooled by heat exchange with the combined hydrocracked feed stream in line 154 and conveyed to a hydrocracked separator 180 for separation. The hydrocracked stream may be separated in the hydrocracked separator 180 to provide a vapor hydrocracked stream in a hydrocracked separator overhead line 182 and a liquid hydrocracked stream comprising a stripper liquid hydrocracked stream in a hydrocracked separator bottom line 184. The hydrocracked separator 180 may be in downstream communication with the hydrocracked reactor 150.

[0096] The hydrocracking separator 180 operates at 177°C (350°F) to 371°C (700°F), and preferably operates at 232°C (450°F) to 315°C (600°F). The hydrocracking separator 180 may operate at a slightly lower pressure than the hydrocracking reactor 150 resulting in a pressure drop through the intervening equipment. The hydrocracking separator 180 may operate at a pressure between 3.4 MPa (gauge) (493 psig) and 20.4 MPa (gauge) (2959 psig). The temperature of the vaporous hydrocracking stream in the hydrocracking separator overhead line 182 may be the operating temperature of the hydrocracking separator 180. The vaporous hydrocracked stream in hydrocracking separator overhead line 182 may be combined with the cooled hydroisomerized stream 52 in hydroisomerization line 50 to provide a mixed hydroisomerized stream in line 54 and separated together in hydroisomerization separator 56. Only the vapor component of the hydrocracked stream is mixed with the liquid component of the hydroisomerized stream to retain heat in the liquid hydrocracked stream in line 184. Thus, the liquid hydrocracked stream is not mixed with the liquid hydroisomerized stream, but remain separated from each other until stripping or fractionation.

[0097] The stripper liquid hydrocracked stream in line 184 may be fed to the hot stripper 102 to strip lighter gases. In an embodiment, the stripper liquid hydrocracked stream in line 184 may be stripped in the hot stripper 102 to remove hydrogen sulfide, naphtha, and lighter gases. The stripper liquid hydrocracked stream in hydrocracking column bottoms line 184 may be fed directly to the hot stripper 102 without undergoing cooling. The hot stripper 102 may be operated at a bottom temperature between 149° C. (300° F.) and 288° C. (550° F.), preferably not exceeding 260° C. (500° F.).

[0098] In an embodiment, the hot stripper 102 may be part of a dual stripping vessel 100 including a cold stripper 86 and a hot stripper 102 separated from each other by an intermediate wall 103. A stripping medium such as steam, which is an inert gas, from a hot stripping medium line 104 may be used to strip light gases from the stripper liquid hydrocracked stream in line 184. The hot stripper 86 provides an overhead stripping stream of naphtha, LPG, hydrogen, hydrogen sulfide, steam and other gases in a hot stripper overhead line 105 and a fractionator liquid hydrocracked stream in a hot stripper bottoms line 106. The hot stripper overhead line 105 is preferably fed to the cold stripper 86 to further strip the lighter boiling components. The fractionator liquid hydrocracked stream in the hot stripper line 106 is preferably heated without a fired heater and fed to the product fractionator 120 at a feed point above the feed point of the fractionator liquid hydroisomerize stream in the cold stripper bottoms line 90. The fractionator liquid hydrocracked stream is fractionated in the product fractionator 120 with the fractionator liquid hydroisomerize stream as previously described.

[0099] It is important to separate the liquid fuel components, particularly the kerosene components, in the hydrocracked stream 158 from the liquid fuel components in the hydroisomerized stream in line 50 to conserve heat in the hydrocracked stream, thereby avoiding the need to heat the fuel components in the hydroisomerized stream back to the kerosene boiling temperature to separate the kerosene components from the diesel components. In one aspect, the hydroisomerized stream in line 50 and the hydrocracked stream in line 158 are completely separated from each other before they are separated in their respective separators 56 and 180, respectively. The liquid hydrocracked stream in line 184 from the hydroisomerized separator 180 is separated from the liquid hydroisomerized stream in line 60 from the hydroisomerization separator 56 and the stripper liquid hydroisomerized stream in line 70 from the cold separator 62. In addition, the fractionator liquid hydroisomerized stream in line 90 is separated from the fractionator liquid hydrocracked stream in line 106.

[0100] Figure 2 yes Figure 1 Another alternative embodiment of the process 10' flashes a liquid hydrocracked stream taken from the hydrocracked stream to provide a fractionator liquid hydrocracked stream separate from stripping the stripper hydroisomerization stream. Figure 1 The same configuration Figure 2 The elements in will have Figure 1 The same reference numerals as in Figure 1 The corresponding components in different configurations Figure 2 Elements in the same figure have the same reference numerals but are represented by a prime ('). Figure 2 The configuration and operation of the embodiment of Figure 1Same as in with the following exceptions.

[0101] Figure 2 The stripping vessel 100' is a single stripping column 86' which receives and strips light gases from the stripper liquid hydroisomerization stream in the cold tower bottoms line 70 with an inert gas, such as steam, from the stripping medium line 89'. The product fractionator 120' provides an overhead gaseous stream of naphtha in an overhead line 122. The fractionator overhead stream may be condensed and separated from water in a fractionation receiver 130'. Unstabilized liquid naphtha from the bottom of the receiver 130' in a fractionator overhead liquid line 132' may be fed to a debutanizer 170 along with a stripper liquid overhead stream in the stripper receiver bottoms line 96 and a rich absorbent stream in a rich absorber bottoms line 166, while a condensed reflux stream is refluxed to the column. The fractionator off-gas stream in line 131 may be compressed in an off-gas compressor 133 to provide a compressed off-gas stream in line 92, and the fractionator off-gas stream is combined with a cold stripper overhead stream in a cold stripper overhead stream line 87' to provide a condenser stream in line 94. The overhead stripping stream in overhead line 87' may be combined with the off-gas compressed stream in line 92 to produce a condenser stream in line 94. The condenser stream in line 94 is condensed and separated by cooling in a stripping receiver 95. The fractionator liquid hydroisomerization stream in stripper bottoms line 90' ​​is fed to a product fractionation column 120' for fractionation.

[0102] The liquid hydrocracked stream in hydrocracking column bottoms line 184' is fed to hydrocracking flash drum 102' to provide a flashed vapor hydrocracked stream in flash column overhead line 105' and a fractionator liquid hydrocracked stream in flash column bottoms line 106'. The vapor flashed hydrocracked stream in flash column overhead line 105' may be fed to stripping column 86' to strip volatiles with the stripped liquid hydroisomerize stream from cold column bottoms line 70. The fractionator liquid hydrocracked stream in flash column bottoms line 106' may be fed to product fractionation column 120 at a location above the feed of the fractionator liquid hydroisomerize stream in line 90'.

[0103] Figure 3 yes Figure 1 Another alternative embodiment of the process 10″ is to strip the cold stripper liquid hydroisomerization stream in line 70 and the stripper liquid hydrocracked stream in line 184 in the same stripper 102″ in a single stripping vessel 100′. Figure 1 The same configuration Figure 3 The elements in will have Figure 1 The same reference numerals as in Figure 1 The corresponding components in different configurations Figure 3The elements in the drawings will have the same reference numerals but will be indicated by double primes ("). Figure 3 The configuration and operation of the embodiment of Figure 1 178. The same as in , with the following exceptions. A single stripped fractionated liquid stream comprising a stripped liquid hydroisomerization stream and a stripped liquid hydrocracked stream in stripping line 106" is fed to product fractionator 120. Fractionator overhead gaseous stream 122 is completely condensed so that fractionator receiver 130 produces only a liquid bottoms stream, a portion of which is refluxed, and the remaining liquid naphtha in overhead liquid line 132 is fed to combine with the naphtha stream in debutanized naphtha stream in line 176 to produce a product naphtha stream in line 178.

[0104] Figure 4 yes Figure 1 Another alternative embodiment of process 10+ wherein the cold liquid hydroisomerization stream in line 70+ and the liquid hydrocracked stream in line 184+ are not stripped but fed directly to product fractionation column 120. Figure 1 The same configuration Figure 4 The elements in will have Figure 1 The same reference numerals as in Figure 1 The corresponding components in different configurations Figure 4 Elements in the same reference numerals will have the same reference numerals but will be indicated by a plus sign (+). Figure 4 The configuration and operation of the embodiment of Figure 1 The same as in, with these foregoing and following exceptions.

[0105] Figure 1 The stripping tower 100 and all its accessories are Figure 4 The fractionator off-gas stream in the fractionator receiver overhead line 131+ is compressed in the off-gas compressor 133+ to provide an off-gas compressed stream in line 92+ and fed to the off-gas scrubber 140+. In addition, the rich absorber bottoms stream in the rich absorber bottoms line 166+ from the bottom of the sponge absorber tower 160 can be fed to the debutanizer 170. The unstabilized liquid naphtha in the fractionator overhead liquid line 132+ from the bottom of the receiver 130+ can be fed to the debutanizer 170, possibly the rich absorbent stream in the rich absorber bottoms line 166+, while the condensed reflux stream is refluxed to the tower.

[0106] Prior to entering the stripping column 100 or the product fractionation column 120, any liquid hydrocracked stream in line 184 or 106 may be removed from Figure 11, 2 or 3. The kerosene in the liquid hydrocracked stream is separated from any liquid hydroisomerized stream in lines 60, 70 and 90 in the product fractionator 120. This arrangement prevents the kerosene in the liquid hydrocracked stream from being cooled and reboiled in the product fractionator 120, which wastes heat. It is contemplated that the liquid hydrocracked stream and the liquid hydroisomerized stream may enter the product fractionator together.

[0107] Example

[0108] The disclosed process was simulated using dedicated separators for the hydroisomerization effluent and the hydrocracking effluent. The duty of each heater or exchanger is provided in the following table. In the table, each heater or exchanger is provided with the reference numeral of the element in the figure, except for the heaters or exchangers not shown in the figure.

[0109] surface

[0110]

[0111]

[0112] It is apparent that the disclosed process significantly reduces the duty in the fractionator reboiler 121 and in the hydroisomerate exchanger 52.

[0113] Specific implementation plan

[0114] While the following is described in conjunction with specific embodiments, it should be understood that this description is intended to illustrate and not to limit the scope of the foregoing description and the appended claims.

[0115] A first embodiment of a method for hydrotreating a hydrocarbon stream includes hydroisomerizing a hydroisomerization feed stream over a hydroisomerization catalyst in the presence of hydrogen to provide a hydroisomerization feed stream; hydrocracking a hydrocracking feed stream in the presence of hydrogen to provide a hydrocracking feed stream; separating a liquid hydroisomerization feed stream from the hydroisomerization feed stream; separating a liquid hydrocracking feed stream from the hydrocracking feed stream separated from the liquid hydroisomerization feed stream; and feeding the liquid hydroisomerization feed stream to a product fractionation tower; and feeding the liquid hydrocracking feed stream to a product fractionation tower. An embodiment of the present disclosure is one, any one, or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, and also includes separating the hydrocracking feed stream in a hydrocracking separator to provide a vapor hydrocracking feed stream and a liquid hydrocracking feed stream, and feeding the liquid hydrocracking feed stream to a fractionation tower. The embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, and also include separating the hydroisomerization stream in the hydroisomerization separator to provide a vapor hydroisomerization stream and a liquid hydroisomerization stream, and feeding the liquid stream to a fractionation tower. The embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the hydroisomerization feed stream is a biorenewable stream. The embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, wherein the hydrocracking feed stream is taken from a product fractionation tower. The embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, and also include cooling the liquid hydroisomerization stream, and separating the cooled liquid hydroisomerization stream into a cold vapor hydroisomerization stream and a cold liquid hydroisomerization stream, and feeding the cold liquid hydroisomerization stream to a fractionation tower. Embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, and further include mixing a vapor hydrocracking stream with a liquid hydroisomerization stream. Embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, and further include stripping a hydroisomerization stream to provide a liquid hydroisomerization stream, and stripping a hydrocracking stream separately from the stripped hydroisomerization stream to provide a liquid hydrocracking stream. Embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the first embodiment in this paragraph, and further include stripping a hydroisomerization stream, and flashing a hydrocracking stream separately from the stripped hydroisomerization stream to provide a liquid hydrocracking stream.Embodiments of the present disclosure are one, any or all of the prior embodiments in this paragraph to the first embodiment in this paragraph, and further include hydrotreating a hydrocarbon stream to provide a hydrotreated stream, and separating the hydrotreated stream to provide a vapor hydrotreated stream and a liquid hydrotreated stream, and removing a hydroisomerization feed stream from the liquid hydrotreated stream.

[0116] A second embodiment is a process for hydrotreating a hydrocarbon stream, the process comprising hydroisomerizing a hydroisomerization feed stream over a hydroisomerization catalyst in the presence of hydrogen to provide a hydroisomerized stream; separating the hydroisomerized stream in a hydroisomerization separator to provide a vapor hydroisomerized stream and a liquid hydroisomerized stream; withdrawing a fractionator hydroisomerization stream from the liquid hydroisomerized stream; hydrocracking a hydrocracking feed stream in the presence of hydrogen to provide a hydrocracked stream; separating the hydrocracked stream in a hydrocracking separator to provide a vapor hydrocracked stream and a liquid hydrocracked stream; withdrawing a fractionator hydrocracked stream from the liquid hydrocracked stream; feeding the fractionator hydroisomerized stream to a product fractionation column; and feeding the fractionator hydrocracked stream to a product fractionation column. Embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, wherein the hydroisomerization feed stream is a biorenewable stream. Embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, and further include cooling the liquid hydroisomerization stream to provide a cooled liquid hydroisomerization stream, and separating the cooled liquid hydroisomerization stream into a cold vapor hydroisomerization stream and a cold liquid hydroisomerization stream, and removing a fractionator hydroisomerization stream from the cold liquid hydroisomerization stream. Embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, and further include mixing a vapor hydrocracking stream with a liquid hydroisomerization stream. Embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, further comprising steam extracting a stripper hydroisomerization stream from a cold liquid hydroisomerization stream to provide a fractionator hydroisomerization stream, and steam extracting a stripper hydrocracking stream from a liquid hydrocracking stream to provide a fractionator hydrocracking stream. Embodiments of the present disclosure are one, any or all of the previous embodiments in this paragraph to the second embodiment in this paragraph, further comprising steam extracting a stripper hydroisomerization stream from a cold liquid hydroisomerization stream to provide a fractionator hydroisomerization stream, and flashing a liquid hydrocracking stream to provide a fractionator hydrocracking stream. Embodiments of the present disclosure are one, any or all of the prior embodiments in this paragraph to the second embodiment in this paragraph, and also include hydrotreating a hydrocarbon stream to provide a hydrotreated stream, and separating the hydrotreated stream to provide a vapor hydrotreated stream and a liquid hydrotreated stream, and removing a hydroisomerization feed stream from the liquid hydrotreated stream.

[0117] A third embodiment of a method for hydrotreating a biorenewable feed stream includes hydroisomerizing a hydroisomerization feed stream over a hydroisomerization catalyst in the presence of hydrogen to provide a hydroisomerized feed stream; hydrocracking a hydrocracking feed stream in the presence of hydrogen to provide a hydrocracked feed stream; separating a liquid hydroisomerization feed stream from the hydroisomerization feed stream; separating the liquid hydrocracked feed stream from the hydrocracked feed stream separated from the liquid hydroisomerization feed stream; feeding the liquid hydroisomerization feed stream to a product fractionation column; and feeding the liquid hydrocracked feed stream to a product fractionation column. Embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the third embodiment in this paragraph, and also include steam extracting a stripper hydroisomerization stream from a hydroisomerization stream to provide a liquid hydroisomerization stream, and steam extracting a stripper hydrocracking stream from a hydrocracking stream separately from the stripper hydroisomerization stream to provide a liquid hydrocracking stream. Embodiments of the present disclosure are one, any one or all of the previous embodiments in this paragraph to the third embodiment in this paragraph, and also include steam extracting a stripper hydroisomerization stream from a hydroisomerization stream to provide a liquid hydroisomerization stream, and flashing a flashed liquid stream taken from a hydrocracking stream separately from the stripper hydroisomerization stream to provide a liquid hydrocracking stream.

[0118] Although there is no further detailed description, it is believed that those skilled in the art can make the most of the present disclosure by using the foregoing description and can easily determine the essential characteristics of the present disclosure without departing from the spirit and scope of the present invention, and can make various changes and modifications of the present disclosure and adapt it to various usages and conditions. Therefore, the aforementioned preferred specific embodiments should be understood as merely illustrative and not to limit the rest of the present disclosure in any way, and are intended to cover various modifications and equivalent arrangements included within the scope of the appended claims.

[0119] In the foregoing, all temperatures are set forth in degrees Celsius and all parts and percentages are by weight unless otherwise indicated.

Claims

1. A method for hydrotreating a hydrocarbon stream, the method comprising: hydroisomerizing a hydroisomerization feed stream over a hydroisomerization catalyst in the presence of hydrogen to provide a hydroisomerized stream; hydrocracking a hydrocracking feed stream in the presence of hydrogen to provide a hydrocracking stream; separating a liquid hydroisomerization stream from the hydroisomerization stream; separating a liquid hydrocracked stream from the hydrocracked stream separated from the liquid hydroisomerization stream; as well as feeding the liquid hydroisomerization stream to a product fractionation column; and feeding the liquid hydrocracked stream to the product fractionation column.

2. The method according to claim 1, further comprising separating the hydrocracked stream in a hydrocracked separator to provide a vapor hydrocracked stream and a liquid hydrocracked stream, and feeding the liquid hydrocracked stream to the fractionation column.

3. The method of claim 2, further comprising separating the hydroisomerization stream in a hydroisomerization separator to provide a vapor hydroisomerization stream and the liquid hydroisomerization stream, and feeding the liquid stream to the fractionation column.

4. The process of claim 1 wherein the hydroisomerization feedstock stream is a biorenewable stream.

5. The process of claim 1, wherein the hydrocracking feed stream is taken from the product fractionation column.

6. The method of claim 3, further comprising cooling the liquid hydroisomerization stream, and separating the cooled liquid hydroisomerization stream into a cold vapor hydroisomerization stream and a cold liquid hydroisomerization stream, and feeding the cold liquid hydroisomerization stream to the fractionation column.

7. The method of claim 6, further comprising mixing the vaporous hydrocracked stream with the liquid hydroisomerized stream.

8. The process of claim 1 further comprising stripping a hydroisomerization stream to provide the liquid hydroisomerization stream, and stripping a hydrocracked stream separately from stripping the hydroisomerization stream to provide the liquid hydrocracked stream.

9. The process of claim 1 further comprising stripping a hydroisomerized stream, and flashing a hydrocracked stream separately from stripping the hydroisomerized stream to provide the liquid hydrocracked stream.

10. The process of claim 1 further comprising hydrotreating a hydrocarbon stream to provide a hydrotreated stream, and separating the hydrotreated stream to provide a vaporous hydrotreated stream and a liquid hydrotreated stream, and removing the hydroisomerization feed stream from the liquid hydrotreated stream.

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