Process for catalytic conversion of naphtha to light olefins in predominant mode

Abstract

A process for converting naphtha to predominantly ethane, comprising contacting a naphtha stream with a zeolite catalyst comprising a noble metal and hydrogen to produce a light paraffin stream comprising ethane and propane; wherein the catalyst comprises at least 0.005% by weight of a noble metal at a ratio of hydrogen to hydrocarbons of less than 5. In a predominantly propane mode, the catalyst comprises 0.005 wt% noble metal and is operated at a reaction temperature of not more than 450 DEG C. The method can switch between modes according to the expectation of the operator.

Description

[0001] Priority declaration

[0002] This application claims priority to U.S. Application No. 63 / 351,785, filed on June 13, 2022, which is incorporated herein in its entirety. Technical Field

[0003] The field is the conversion of naphtha to light olefins. The field may particularly relate to the conversion of naphtha to light olefins in a two-step conversion process. Background Art

[0004] Light olefin production is critical to producing enough plastics to meet global demand. Paraffin dehydrogenation (PDH) is a process in which light paraffins (such as propane and butane) can be dehydrogenated to produce propylene and butenes, respectively. Dehydrogenation is an endothermic reaction that requires external heat to drive the reaction to completion.

[0005] Fluid catalytic cracking (FCC) is another endothermic process that can be fine-tuned to produce large amounts of propylene. However, not every FCC unit can be fine-tuned to make large amounts of propylene. Moreover, high-propylene FCC units do not make much ethylene; less than 1% of the global ethylene supply comes from FCC.

[0006] A large amount of ethylene consumed in the production of plastics and petrochemicals such as polyethylene is produced by thermal cracking of hydrocarbons. Steam is often mixed with the feed stream to the cracking furnace to reduce the hydrocarbon partial pressure and increase the olefin yield, and to reduce the formation and deposition of carbonaceous materials in the cracking reactor. Therefore, the process is often referred to as steam cracking or pyrolysis. Ethane oxidative dehydrogenation is a newer catalytic process for converting ethane to ethylene that can be performed at lower temperatures with less carbon oxide emissions than steam cracking.

[0007] Two types of feeds are commonly used for steam cracking. Ethane feed is used in areas where light hydrocarbon gases are prevalent. In areas where natural gas is not abundant, naphtha feed is used for steam cracking. Since the production cost is higher than that of pyrolysis ethane cracking, pyrolysis naphtha cracking has long determined the price of the ethylene industry. The world does not currently produce enough ethane to meet the growing demand for ethylene. Therefore, areas such as Asia and Europe where ethane is in short supply rely mainly on naphtha cracking to supply ethylene. Naphtha cracking only produces 30% to 35% of ethylene, and the remainder includes both relatively high-value by-products including propylene, butadiene and butene-1 and relatively low-value by-products including pyrolysis oil, pyrolysis gas and fuel gas. The additional pressure of naphtha cracking (including minimum generation requirements and environmental issues) has led to government disapproval in some regions (such as China). The ethylene industry needs a more efficient, economical and environmentally friendly route for light olefins from naphtha feed. Summary of the invention

[0008] A process for converting naphtha to predominantly ethane in a mode comprising contacting a naphtha stream with a zeolite catalyst comprising a noble metal and hydrogen to produce a light paraffin stream comprising ethane and propane; wherein the catalyst comprises at least 0.005 wt. %, preferably 0.02 wt. % of the noble metal at a hydrogen to hydrocarbon ratio of less than 5. In the predominantly propane mode, the catalyst comprises at least 0.005 wt. % of the noble metal, and preferably 0.01 wt. % to 0.1 wt. % of the noble metal, and is operated at a reaction temperature not exceeding 440° C. The process can be switched between modes as desired by the operator. BRIEF DESCRIPTION OF THE DRAWINGS

[0009] Figure 1 is a schematic diagram of the method and apparatus of the present disclosure.

[0010] Figure 2 yes Figure 1 Schematic diagram of a method and apparatus of an alternative embodiment.

[0011] Figure 3 yes Figure 1 Schematic diagrams of methods and apparatus of additional alternative embodiments.

[0012] Figure 4 yes Figure 1 Schematic diagram of a method and apparatus of another alternative embodiment.

[0013] Figure 5 is a plot of the conversion of a 50 / 50 mixture of isobutane and n-butane via an olefin intermediate.

[0014] definition

[0015] The term "communication" means that fluid flow is operatively allowed between the enumerated components, which can be characterized as "fluid communication".

[0016] The term "downstream communication" means that at least a portion of the fluid flowing to the body in the downstream communication can operatively flow from the object in fluid communication therewith.

[0017] The term "upstream communication" means that at least a portion of the fluid flowing out of the body in the upstream communication can be operatively flowed to the object in fluid communication therewith.

[0018] The term "direct communication" means that the fluid flows from the upstream component into the downstream component without passing through any other intervening container.

[0019] The term "indirect communication" means that the fluid flow from the upstream component enters the downstream component after passing through an intervening container.

[0020] The term "bypass" means that the object loses downstream communication with the bypassed body at least within the scope of the bypass.

[0021] As used herein, the term "predominantly" or "majority" means greater than 50%, suitably greater than 75%, and preferably greater than 90%.

[0022] The term "Cx" is understood to refer to molecules having the number of carbon atoms represented by the subscript "x". Similarly, the term "Cx-" refers to molecules containing less than or equal to x, and preferably x and fewer carbon atoms. The term "Cx+" refers to molecules having greater than or equal to x, and preferably x and more carbon atoms.

[0023] 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.

[0024] 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. DETAILED DESCRIPTION

[0025] In the proposed process, the C3-C8+ hydrocarbon feedstock is first stored and loaded into a "naphtha to ethane and propane" unit to convert the naphtha into the desired ethane and propane and the less desired methane. The produced ethane is fed to an ethylene production unit. These units provide ethane to ethylene yields of over 75%. The produced propane is fed to a propylene production unit, which provides a propane to propylene yield of over 85%. The methane byproduct from the naphtha conversion unit and the ethane and propane production units can be used as fuel, including fuel required to operate the ethylene and propylene production units operating at high temperatures. The unconverted or under-converted C4+ components in the reactor outlet can be recycled to the reactor inlet for further processing into ethane and propane. Aromatics can also be recovered and further processed.

[0026] Go to Figure 1, the naphtha stream in line 10 can be combined with the hydrogen stream in line 22 and the heavy stream in line 12 to provide a charge stream in line 11, heated and fed to the naphtha to ethane and propane (NEP) reactor 16 for contact with the NEP catalyst. The hydrogen stream in line 22 can be supplemented or provided by the ethylene production unit 30 or the propylene production unit 40. The naphtha stream may include C4 to C12 hydrocarbons, preferably having a T10 between -10°C and 60°C and a T90 between 70°C and 180°C. The naphtha feed stream may include normal paraffins, isoparaffins, cycloparaffins and aromatics. The naphtha stream can be heated to a reaction temperature between 300°C and 550°C, and preferably between 325°C and 525°C. The weight space velocity should be 0.5hr -1 With 10hr -1 between, and preferably within 1 hr -1 With 4hr -1 The molar ratio of hydrogen to hydrocarbon should be between 0.5 and 5, and the total pressure should be between 0.1 MPa (absolute) and 2 MPa (absolute). Under these conditions, the C2-C4 yield is always in excess of 80 wt%, while the methane yield is less than 10 wt%, suitably less than 8 wt%, and typically less than 6 wt%, and preferably no more than 5 wt%. In another embodiment, the molar ratio of hydrogen to hydrocarbon may typically be no more than 5, suitably no more than 3, and preferably no more than 2. Under these conditions, ethane may account for more than 60 wt% of the total C2 to C3 and C2 to C4 produced in the NEP reactor 16. In the NEP reactor 16, sufficient hydrogen is co-fed with the hydrocarbon to direct conversion to a product consisting primarily of ethane or propane, while maintaining a hydrogen to hydrocarbon ratio low enough to promote the desired reaction kinetics initiated by dehydrogenation.

[0027] The NEP catalyst dehydrogenates naphtha molecules to their olefin analogs, interconverts olefins to lighter olefins and hydrogenates the lighter olefins to produce a light paraffin stream including ethane and propane. Interconversion may mean that olefins are also oligomerized to higher carbon olefins, and then these higher carbon olefins are cracked to lower carbon olefins. This chemical mechanism avoids or minimizes hydrogenolysis and pyrolytic cracking reactions that produce methane. Methane is an undesirable byproduct that represents an opportunity for the production of valuable ethane and propane, and consumes excess hydrogen.

[0028] The NEP catalyst for converting naphtha into ethane and propane may comprise a molecular sieve, including macroporous or mesoporous molecular sieves, that is, including 10-membered or 12-membered rings, respectively. Examples of suitable molecular sieves include MFI, MEL, MFI / MEL co-biopolymers, MTW, TUN, UZM-39, IMF, UZM-44, UZM-54, MWW, UZM-37, UZM-8, UZM-8HS. Examples of suitable molecular sieves also include FER, AHT, AEL (SAPO-11), AFO (SAPO-41), MRE, MFS, EUO-1, TON (ZSM-22), MTT (ZSM-23), and UZM-53. Additional molecular sieves with larger pores include FAU, EMT, FAU / EMT co-biopolymers, UZM-14, MOR, BEA, UZM-50, MTW, ZSM-12. Additional examples include MSE and UZM-35.

[0029] MFI is a suitable NEP catalyst. It should be understood that ZSM-5 is an aluminosilicate zeolite belonging to the MFI type of the pentasil zeolite family and has a chemical formula of Na n A1 n Si 96-n O 92 .16H2O (0 < n < 10). In various embodiments, the ZSM-5 zeolite may have a silica / alumina molar ratio of 20 to 1000, 20 to 800, 20 to 600, 25 to 400, 25 to 200, or 25 to 80. In various embodiments, the ZSM-5 zeolite may have a crystal size in the range of 10 nm to 600 nm, 20 nm to 500 nm, 30 nm to 450 nm, 40 nm to 400 nm, or 50 nm to 300 nm.

[0030] The NEP catalyst may include a bound zeolite. The binder may include oxides of aluminum, siloxane, zinc, titanium, zirconium, and mixtures thereof. In an embodiment, the binder may include phosphate in the binder or phosphate of the aforementioned oxide binder materials. Preferably, the binder is aluminum phosphate. The MFI zeolite may be loaded in a binder containing alumina, such as aluminum phosphate.

[0031] The MFI zeolite slurry may first be mixed with a binder in the form of a colloidal suspension (sol) and a gelling agent, and then dropped into hot oil to form controlled spheres to produce a calcined carrier with a diameter of 1 / 64 inch to 1 / 10 inch. The spheres can be washed with ammonia to remove sodium ions from the zeolite, dried, and calcined to remove the organic structure-directing agent (OSDA) from the synthesized zeolite. Optionally, the calcined carrier can be subjected to ammonium ion exchange using an ammonium nitrate solution to remove residual sodium ions and dried at 110 °C.

[0032] In an embodiment, the NEP catalyst includes a metal on the catalyst. The metal may include a transition metal. In another embodiment, the metal may include platinum, palladium, iridium, rhenium, ruthenium, and mixtures thereof. The metal may be a noble metal. In additional embodiments, a modifier metal may also be included on the catalyst. The modifier metal may include tin, germanium, gallium, indium, thallium, zinc, silver, and mixtures thereof. The modifier metal should be more concentrated on the binder than on the zeolite. In an embodiment, 0.01 wt % to 5 wt % of each of the transition metal and the modifier metal is on the catalyst.

[0033] The metal may be incorporated into the binder by evaporative impregnation.A platinum solution (such as tetraamine nitrate platinate or chloroplatinic acid) may be contacted with the spherical support which has been calcined and ion exchanged in a rotary evaporator followed by drying and oxidation.

[0034] In a preferred embodiment, the NEP catalyst comprises a metal on a spherical support of the catalyst. Preferably, there is more metal on the binder than on the zeolite. In an embodiment, at least 60 wt%, suitably at least 70 wt%, preferably at least 80 wt% and most preferably at least 90 wt% of the metal is on the binder. In an embodiment, the zeolite and / or the entire NEP catalyst is steam oxidized to remove the metal from the zeolite. Steaming is preferably performed after the metal is added to the catalyst. The dried metal-impregnated binder support can be steam oxidized in air for a sufficient time to provide a NEP catalyst. Steam oxidation in air for about 1 hour to 3 hours at a temperature of 500° C. to 650° C. and 10 mol % to 30 mol % steam may be suitable.

[0035] The NEP catalysts are preferably reduced to activate them to catalyze the NEP reaction. For example, the catalysts may be reduced in flowing hydrogen at 500°C to 550°C for 2 hours to 5 hours prior to contacting the feed.

[0036] After paraffin conversion, a light paraffin stream is discharged from the NEP reactor 16 into an effluent line 18. The light paraffin stream may include at least 40 wt% ethane or at least 40 wt% propane or at least 70 wt% and preferably at least 80 wt% ethane and propane. The ratio of ethane to propane may be in the range of 0.1 to 5. The light paraffin stream may have less than 15 wt%, suitably less than 12 wt%, more suitably less than 10 wt%, preferably less than 8 wt%, more preferably less than 6 wt%, and most preferably less than 5 wt% methane.

[0037] The ratio of ethane to propane (E / P) in the light paraffin stream in line 18 from the NEP reactor 16 can be varied in the presence of a noble metal-containing catalyst by simply adjusting the operating temperature. For example, at low metal concentrations such as 0.01 wt % to 0.1 wt % noble metal, an E / P of 2:1 to 1:2 can be achieved at relatively high temperatures of greater than 440° C. In addition, at the same noble metal concentration, a propane to ethane ratio (P / E) of up to 5:1 to 3:1 with a propane yield of not less than 60% can be obtained at relatively low temperatures of less than 420° C. Further fine-tuning of ethane and propane production outputs can be achieved by adjusting the hydrogen to hydrocarbon ratio, total pressure and / or WHSV, depending on the supply and demand of ethylene and propylene. At high metal concentrations such as 0.2 wt % to 1.0 wt % noble metal, an E / P of 1:1 to 7:1 can be achieved.

[0038] It is found that from 1.0 to greater than 2 is unfavorable to ethane production and favorable to propane production. To achieve a given ethane content, a lower temperature is required at a lower hydrogen to hydrocarbon ratio while suppressing methane production. At temperatures of 400°C or less, a reduced hydrogen to hydrocarbon ratio such as 1.0 to 2.0 favors propane production.

[0039] The presence of platinum is most critical to obtain a respectable ethane yield of greater than 10%. At a platinum content as low as 50wppm, high ethane of greater than 40% can be obtained. Note that increasing the platinum content beyond 0.1 wt% increases the C2 formation level to greater than 60% while maintaining a low methane formation level of less than 6% to 8%.

[0040] The process can be operated in a mode that is mainly ethane, which includes at least 0.005% by weight of precious metal on the catalyst, and a ratio of hydrogen to hydrocarbons of less than 5, and suitably no more than 2, and preferably no more than 1.5. The catalyst may contain less than 0.5% by weight of precious metals, but may contain at least 0.02% by weight, and preferably at least 0.1% by weight of precious metals. At a temperature of at least 425°C, lower pressures and lower ratios of hydrogen to hydrocarbons are better for ethane production, so the reaction pressure in terms of reactor outlet pressure can be less than 20 bar absolute pressure, and preferably less than 15 bar absolute pressure. Higher temperatures are better for ethane production, so the reaction temperature is suitably greater than 440°C, more suitably greater than 450°C, and preferably greater than 460°C. The precious metal is typically platinum. Under these conditions of the mode that is mainly ethane, the E / P ratio can be in the range of greater than 1:1 to 7:1.

[0041] The process can also be operated in a predominantly propane mode, which includes a noble metal concentration of 0.01 wt % to 0.1 wt %, and preferably greater than 0.02 wt %. Lower temperatures such as no more than 440° C. are suitable, and preferably no more than 425° C. for predominantly propane production. In the propane mode, the ratio of hydrogen to hydrocarbons and the total pressure will depend on the operating temperature and WHSV. At a temperature of 400° C. or less and a WHSV of 2.0 or less, operation at a ratio of hydrogen to hydrocarbons of not less than 5.0 and more preferably not less than 2.0 is conducive to propane production of 50% or more yield. In the predominantly propane mode, the minimum temperature should be 350° C. In the predominantly propane mode, the ratio of hydrogen to hydrocarbons should be at least 2, suitably at least 2.5, and preferably at least 3. The total pressure for propane production should be less than 25 bar, and preferably less than 20 bar. Using a precious metal catalyst, a predominantly propane mode can achieve P / E ratios of greater than 1:1 to 5:1 while still producing over 60 wt% propane yields while producing over 75 wt% combined ethane and propane total yields.

[0042] The operator may prefer to change the operating conditions from a predominantly ethane mode to a predominantly propane mode or vice versa with or without changing the catalyst. Adjustments will include reaction pressure and temperature changes and hydrogen to hydrocarbon ratios. Adjusting the output of ethane and propane can also respond to fluctuating market demands.

[0043] For example, the operator may switch from operating in a predominantly ethane mode with a hydrogen to hydrocarbon ratio of less than 2 to a predominantly propane mode by increasing the hydrogen to hydrocarbon ratio to at least 2, and suitably at least 2.5. Alternatively, the operator may switch from operating in a predominantly propane mode with a hydrogen to hydrocarbon ratio of at least 2, or suitably at least 2.5, to a predominantly ethane mode by decreasing the hydrogen to hydrocarbon ratio to less than 2.

[0044] The operator may switch from operating in a predominantly propane mode at a reaction temperature not exceeding 440° C. to operating in a predominantly ethane mode at a reaction temperature exceeding 440° C. Alternatively, the operator may switch from operating in a predominantly ethane mode at a reaction temperature exceeding 440° C. to operating in a predominantly propane mode at a reaction temperature not exceeding 440° C.

[0045] When switching to the main propane mode, the operator can increase the reaction pressure from less than 7 bar in the main ethane mode to not less than 7 bar. Alternatively, when switching to the main ethane mode, the operator can reduce the reaction pressure from not less than 7 bar in the main propane mode to less than 7 bar.

[0046] When operating at a given pressure and space velocity, reducing the ratio of hydrogen to hydrocarbons will increase ethane and decrease methane at 450° C. and 465° C., respectively. By adjusting the ratio of hydrogen to hydrocarbons, light paraffin production can be switched to primarily ethane, a balance of ethane and propane, or primarily propane. When operating at a given hydrogen to hydrocarbon ratio and pressure, the space velocity can be adjusted to switch light paraffin production from primarily ethane at a lower space velocity to primarily propane at a higher space velocity, or vice versa.

[0047] When operating with a zeolite catalyst containing at least 0.1 wt. % of a noble metal, the operator can achieve a catalyst temperature of less than 4 hours. -1 The WHSV is switched to the main ethane mode at more than 4hr -1 The main propane mode under the WHSV. Under the same conditions, the operator can go from more than 4 hours -1 The WHSV switches to the main propane mode in less than 4 hours -1 Commercial operation at a flexible WHSV can be easily implemented by adjusting the number of reactors put into production.

[0048] Operating the conversion of naphtha to ethane and propane with a zeolite catalyst comprising at least 0.005 wt. % noble metal and more suitably 0.01 wt. % to 0.2 wt. % noble metal may allow switching between a predominantly ethane production mode and a predominantly propane production mode without changing the catalyst.

[0049] On the other hand, the method contemplates switching the catalyst to a switching mode. For example, when switching from a primarily propane mode to a primarily ethane mode, the operator may unload a catalyst having less than 0.1 wt % of precious metals and load a catalyst having at least 0.1 wt % of precious metals. Alternatively, when switching from a primarily ethane mode to a primarily propane mode, the operator may unload a catalyst having at least 0.1 wt % of precious metals and load a catalyst having less than 0.1 wt % of precious metals.

[0050] It was found that the presence of sulfur in the NEP reaction does not significantly affect conversion. At least 200 wppm of sulfur can be present in the feed without significant effect on conversion. It is expected that the NEP catalyst can handle up to 400 wppm and possibly 500 wppm of sulfur without significant effect on conversion.

[0051] The light paraffin stream can be cooled and fed to the NEP separation unit 20. The NEP separation unit 20 can be a fractionation tower or a series of fractionation towers and other separation units that can separate the light paraffin stream in line 18 into a hydrogen stream in line 22, an ethane stream in line 24, a propane stream in line 26, and a heavy stream in line 12. The NEP separation unit 20 can include a demethanizer that separates the light paraffin stream into a gas stream in the tower top line and a C2+ paraffin stream in the tower bottom line. The gas stream can be sent to a hydrogen purification unit such as a PSA unit to recover the hydrogen in line 22 for recycling to the NEP reactor 16. The remaining methane from the hydrogen purification unit can be used for fuel gas. The C2+ paraffin stream can then be fed to a deethanizer to produce an ethane stream in the deethanizer top line 24 and a C3+ paraffin stream in the deethanizer bottom line. The C3+ paraffin stream may then be fed to a depropanizer to produce a propane stream in depropanizer overhead line 26 and a heavy paraffin stream that may include C4+ hydrocarbons in recovery line 12. The NEP separation unit 20 may take other forms.

[0052] For example, the NEP separation unit 20 may omit the demethanizer, and the light paraffin stream in line 18 may be fed to a deethanizer that produces a C2 stream in a deethanizer overhead line. The C2 stream may be separated in a hydrogen purification unit to recover a hydrogen stream in line 22, and the residual ethane and methane from the hydrogen purification unit may include or supplement the ethane stream in line 24. The hydrogen purification unit may include a membrane unit, and the hydrogen recovered from the membrane unit may be further purified in an absorber and then recovered in line 22 to the NEP reactor 16. In an additional alternative, the C2 stream from the deethanizer may be charged to an ethylene production unit 30, where ethane is converted to ethylene, but methane and hydrogen are passed inertly to be recovered in a downstream ethylene recovery unit.

[0053] The ethane stream in line 24 can be loaded into an ethylene production unit 30, in which the ethane in the ethane stream is converted into ethylene. In an embodiment, the ethylene production unit 30 is a steam cracking unit. The ethane stream in line 24 can be cracked under steam in a furnace to produce a cracked stream including an ethylene stream 32. The ethane stream can be loaded into the ethane steam cracking unit in a gas phase. The ethane steam cracking unit can preferably be operated at a temperature of 750°C (1382°F) to 950°C (1742°F). The cracked stream of the furnace leaving the ethane steam cracking unit can be in a superheated state. One or more quench towers or other devices known in the art, but preferably an oil quench tower and / or a water quench tower, can be used to quench or separate the cracked stream into multiple cracked streams. The ethane steam cracking unit may also include additional distillation towers, amine scrubbers, compressors, expanders, etc., to separate the cracked stream into a cracked stream rich in individual light olefins, wherein the main component is the ethylene stream in line 32. The ethylene stream may include an ethylene yield of at least 75 wt%, preferably at least 80 wt%, based on the ethane stream in line 24. Among other components in the cracked stream leaving the ethane steam cracking, ethylene production unit 30 may be hydrogen, methane, propylene, butenes, and pyrolysis gases. Each of these components may be recovered and further processed.

[0054] The ethylene stream in line 32 and the propylene stream from the ethylene production unit 30 can be recovered or sent to a polymerization unit, a chemical plant or exported. The butene stream can be recovered and used to produce plastics or other petrochemical products by a process such as polymerization, or exported. Based on the ethane stream in line 32, product recovery of at least 50 weight percent, typically at least 60 weight percent and suitably at least 70 weight percent of valuable ethylene, propylene and butene products can be achieved from the ethane steam cracking unit 30. The recovered hydrogen can be exported or recycled to the NEP reactor 16, possibly as a supplement to line 22.

[0055] In another embodiment, ethylene production unit 30 can be an oxidative dehydrogenation (ODH) unit. The ethane stream in pipeline 24 can be loaded into the ODH unit. The ethane ODH process is an alternative technology for ethane steam cracking or ethane pyrolysis for converting ethane into ethylene. Ethane ODH involves contacting ethane feed and an oxygen source under the conditions of at least a portion of oxidative dehydrogenation of ethane in the presence of an ODH catalyst in an ODH reaction zone to produce a product stream, including ethylene, carbon oxides, water and unreacted oxygen, acetic acid and other organic acids and unconverted ethane. The oxygen source can be an oxygen-containing stream or an oxygen-containing material, such as a metal oxide. It has been found that mixed metal oxide catalysts are effective for oxidative dehydrogenation reactions.

[0056] The ODH reactor may use a mixed metal oxide catalyst and operate at a temperature of 300° C. to 400° C. and produce more than 90% ethylene and acetic acid, both of which are useful products. The ODH unit produces an ethylene stream and an acetic acid stream in line 32. Ethane oxidative dehydrogenation using a mixed metal oxide catalyst may be carried out at a temperature of 300° C. to 500° C., preferably 350° C. to 450° C., at a pressure of 0.1 bar to 20 bar, preferably 0.1 bar to 10 bar, 1000 cm 3 / (gcat·hr) to 5000cm 3 The ODH of ethane is carried out at a space velocity of 1.5:1 to 2:1 with a molar ratio of ethane to oxygen and sufficient inert diluent to achieve safe operating conditions. MoVNbTe oxide catalysts (and other related materials) with M1-type structure are believed to be the best catalysts for the ODH of ethane.

[0057] Ethane ODH using a mixed metal oxide catalyst (where the catalyst is also the oxygen source) can be carried out at a temperature of 600° C. to 900° C., preferably 750° C. to 850° C., at a pressure of 0.1 bar to 20 bar, preferably 0.1 bar to 10 bar, and for 1000 h. -1 Up to 5000h -1 Mg6MnO8 oxide catalysts (and other related materials) have been identified as preferred materials for ethane ODH processes where the catalyst is also the oxygen source.

[0058] The propane stream in line 26 can be loaded into a propylene production unit 40, where the propane in the propane stream is converted into propylene. The propylene production unit 40 can be a paraffin dehydrogenation (PDH) unit. The PDH catalyst is used in a dehydrogenation reaction process to catalyze the dehydrogenation of paraffins such as propane. The conditions in the dehydrogenation reactor can include a temperature of 500° C. to 800° C., a pressure of 40 kPa to 310 kPa, and a catalyst to oil ratio of 5 to 100.

[0059] The dehydrogenation reaction can be carried out in a fluidized manner so that a gas with or without a fluidizing inert gas, which may contain the reactant alkane, is distributed to the reactor in a manner that lifts the dehydrogenation catalyst in the reactor vessel while catalyzing the dehydrogenation of the alkane. During the catalytic dehydrogenation reaction, coke is deposited on the dehydrogenation catalyst, resulting in reduced catalyst activity. The dehydrogenation catalyst must then be regenerated in a regenerator. The regenerator can burn coke from the dehydrogenation catalyst and the fuel gas to ensure that there is enough enthalpy in the dehydrogenation reactor to promote the endothermic reaction.

[0060] The selected dehydrogenation catalyst should minimize cracking reactions and facilitate dehydrogenation reactions. Suitable catalysts for use herein include active metals that can be dispersed in porous inorganic support materials such as silica, alumina, silica-alumina, zirconia, or clay. Exemplary embodiments of the catalyst include alumina or silica-alumina containing gallium, noble metals, and alkali or alkaline earth metals.

[0061] The catalyst support comprises a support material, a binder and an optional filler material to provide physical strength and integrity. The support material may include alumina or silica-alumina. Silica sol or alumina sol may be used as a binder. Alumina or silica-alumina typically contains alumina in γ, θ and / or δ phases. The nominal diameter of the catalyst support particles may be 400 microns to 5000 microns, with an average diameter of 600 microns to 3500 microns. Preferably, the surface area of ​​the catalyst support is 85 m 2 / g to 140m 2 / g.

[0062] The fluidized dehydrogenation catalyst may include a dehydrogenation metal on a support. The dehydrogenation metal may be one or a combination of transition metals. A noble metal may be a preferred dehydrogenation metal, such as platinum or palladium. Gallium is an effective metal for paraffin dehydrogenation. The metal may be deposited on the catalyst support by impregnation or other suitable methods, or may be included in a support material or binder during catalyst preparation.

[0063] The acid function of the catalyst should be minimized to prevent cracking and facilitate dehydrogenation. Alkali metals and alkaline earth metals may also be included in the catalyst to reduce the acidity of the catalyst. Rare earth metals may be included in the catalyst to control the activity of the catalyst. Metals in a concentration of 0.001 wt % to 10 wt % may be incorporated into the dehydrogenation catalyst. In the case of noble metals, it is preferred to use 10 parts per million (ppm) by weight to 600 ppm by weight of noble metals. More preferably, it is preferred to use 10 ppm to 100 ppm by weight of noble metals. The preferred noble metal is platinum. Gallium should be present in a range of 0.3 wt % to 3 wt %, preferably 0.5 wt % to 2 wt %. Alkali metals and alkaline earth metals may be present in a range of 0.05 wt % to 1 wt %.

[0064] The regenerated catalyst can be contacted with the propane stream in line 26, perhaps with fluidizing gas, to lift the propane stream and dehydrogenation catalyst into the riser when dehydrogenation occurs. Above the riser, the spent dehydrogenation catalyst and the propylene product can be separated by a centripetal separation device. The propylene product gas can be quenched with a cooling fluid to prevent overreaction from generating undesirable by-products. The separation of the propylene product can include quenching contact and fractionation to produce a propylene product stream in line 42. Unreacted propane can be recycled to the dehydrogenation reactor, and light gases can be recycled to the regenerator as a fuel gas to burn, thereby providing enthalpy for the reaction.

[0065] The propylene production unit may also employ a catalytic moving bed reactor. The reactor section may include several radial flow reactors in parallel or in series that are heated by charge and interstage heaters. A propane stream to which hydrogen may be added flows from a screened center tube through an annular dehydrogenation catalyst bed to an external effluent ring in each dehydrogenation reactor. The flow may be in the opposite manner. The dehydrogenation catalyst may include a noble metal and a mixture thereof, a modifier selected from the group consisting of an alkali metal or alkaline earth metal and their combination, a component selected from the group consisting of tin, germanium, lead, indium, gallium, thallium and their combination, and a porous carrier that forms catalyst particles. The catalyst carrier may include an oil droplet alumina sphere.

[0066] The dehydrogenation conditions may include a temperature of 400° C. to 900° C., a pressure of 0.01 atmospheres absolute to 10 atmospheres absolute, and a reaction temperature of 0.1 hr. -1 Up to 100hr -1 The pressure in the dehydrogenation reactor is maintained as low as practicable, consistent with equipment limitations, to maximize chemical equilibrium advantages. The spent dehydrogenation catalyst in the annular catalyst bed can be removed from the bottom of the bed and transferred to a regenerator to burn the coke from the catalyst at 450°C to 600°C in air. The precious metal on the catalyst can be redispersed by an oxyhalogenation process, dried and returned to the top of the dehydrogenation catalyst bed as regenerated dehydrogenation catalyst.

[0067] The dehydrogenation effluent from the propylene production unit 40 may be cooled, compressed, dried, and hydrogen cryogenically separated from the hydrocarbons, the net gas purity of the hydrogen being 85 mol% to 93 mol%. The hydrocarbon liquids are selectively hydrogenated to convert diolefins and acetylene, and the hydrocarbon liquids are fractionated in a deethanizer to remove ethane, and propylene is separated from propane in a propane-propylene splitter to provide polymer grade propylene in line 42. The propane may be recycled as feed to the propylene production unit 40.

[0068] The heavy stream in line 12 that can be taken from the bottom of the depropanizer can include C4+ paraffins. The heavy stream in line 12 can be recycled to the NEP reactor 16 by combining with the paraffin stream in line 10 and the hydrogen stream in line 22 and charged to the NEP reactor in line 11 to produce more ethane and propane.

[0069] In an alternative embodiment, the light paraffin stream in line 18 may be separated by the NEP separation unit 20 into a hydrogen stream in line 22, an ethane stream in line 24, and a heavy stream in line 12, as previously described, but the propane stream in line 26 may include isobutane. In this embodiment, the propane and isobutane stream in line 26 may be fed to a propylene production unit 40. The propylene production unit 40 may be equipped to dehydrogenate the propane in the propane and isobutane stream 26 to propylene, and to dehydrogenate the isobutane in the propane and isobutane stream in line 26 to isobutylene in the same dehydrogenation reactor. The fractionation section from the propylene production unit 40 may include a depropanizer downstream of the bottoms line of the propylene-propane splitter to separate the unreacted propane overhead from the C4 hydrocarbons. An isobutylene-isobutane splitter in downstream communication with the bottom of the depropanizer may provide isobutylene at the top of the column that may be recovered as a product, and unreacted isobutane at the bottom of the splitter may be returned to the propylene production unit 40 to be converted into isobutylene.

[0070] Unreacted n-C4+ hydrocarbons including n-C4 and higher paraffins in heavy stream 12 can be recovered to be combined with the naphtha stream in feed line 10 and the hydrogen stream in line 22 and charged to the NEP reactor 16 in line 11.

[0071] Figure 2 Shows Figure 1 An alternative embodiment of the embodiment of the invention employs an embodiment of the n-butene conversion unit 50. 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 1 Same as in with the following exceptions.

[0072] In addition to the hydrogen stream in line 22, the ethane stream in line 24, and the propane and isobutane streams in line 26, the NEP separation unit 20' additionally provides an n-butane product stream in line 46. The n-butane stream in line 46 can be fed to a butene production unit 50 to convert the n-butane in the n-butane stream into n-butenes. This embodiment may be useful in situations where a propylene production unit cannot process propane and n-butane in the same dehydrogenation reactor. The butene production unit 50 may include, for example, a propylene production unit for Figure 1 The dehydrogenation unit is described in the propylene production unit of the embodiment.

[0073] The butene production unit 50 can feed the normal butene stream to the normal butene-normal butane separation tower. The normal butene stream in the separator tower top line can be regarded as the normal butene product in the pipeline 52, while the unreacted normal butane in the tower bottom line can be recycled back to the dehydrogenation unit in the butene production unit 50 to be converted into normal butene.

[0074] In some cases, butene production unit 50 may be a butene production unit capable of converting both normal butane and isobutane into normal butene and isobutene. In this case, line 46 carries a propane-depleted butane stream to butene production unit 50, and the propane stream in line 26 carries a butane-depleted propane stream to the propylene production unit. The butene production unit will produce a butene product stream in line 52 that may include isobutene and butene.

[0075] The heavy stream in line 12 from the NEP separation unit 20' may include a C5+ hydrocarbon stream, but it may include C4 hydrocarbons and include a C4+ hydrocarbon stream. Unreacted C4+ or C5+ hydrocarbons in the heavy stream 12 including C4 or C5 and heavier hydrocarbons may be recovered to be combined with the naphtha stream in the feed line 10 and the hydrogen stream in the line 22 and charged to the NEP reactor 16 in the line 11.

[0076] It is also found that the heavy paraffin stream in line 12 may also include aromatics, benzene, toluene, and xylenes. Figure 3 Shows Figure 1 An alternative embodiment of the ... 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 3 The 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 Same as in with the following exceptions.

[0077] The heavy stream in line 12" comprising C4+ paraffins and aromatics, benzene, toluene and xylenes is mixed with the hydrogen stream in line 64, heated and charged to the hydroprocessing reactor 60.

[0078] The hydroprocessing reactor 60 may have one or more hydroprocessing catalyst beds to saturate the aromatic rings in the heavy stream. The heavy stream may be charged to the hydroprocessing reactor 60, and the hydroprocessing inlet temperature may be in the range of 200°C (392°F) to 400°C (752°F). If more than one catalyst bed is used, the hydroprocessing reactor 60 may employ an interlayer hydrogen quench flow.

[0079] Suitable hydroprocessing catalysts are any known conventional hydroprocessing catalysts, and include those consisting of at least one Group VIII metal (preferably iron, cobalt and nickel, more preferably cobalt and / or nickel) and at least one Group VI metal (preferably molybdenum and tungsten) on a high surface area support material (preferably alumina). Other suitable hydroprocessing catalysts include zeolite catalysts, and noble metal catalysts, wherein the noble metal is selected from palladium and platinum. It is within the scope of the present specification to use more than one type of hydroprocessing catalyst in the same hydroprocessing reactor 60. Group VIII metals are typically present in an amount in the range of 2 wt % to 20 wt %, preferably 4 wt % to 12 wt %. Group VI metals will typically be present in an amount in the range of 1 wt % to 25 wt %, preferably 2 wt % to 25 wt %. Typically, hydroprocessing conditions include pressures of 700 kPa (100 psig) to 21 MPa (3000 psig). The hydroprocessing outlet temperature may be in the range of between 300 ° C (572 ° F) and 427 ° C (800 ° F).

[0080] The saturated aromatics and C4+ paraffins in the hydrotreated heavy stream in line 62 can be recovered to be added to the naphtha stream in line 10 and the hydrogen stream in line 22 and charged to the NEP reactor 16 in line 11. Hydrogen from the hydrotreating reactor 60 can be recycled to the NEP reactor 16 along with the saturated aromatics to reduce or eliminate the hydrogen demand from line 22.

[0081] In another embodiment, NEP separation unit 20+ comprises a debutanizer that separates C4 and possibly C5 hydrocarbons for recycling in line 12+ to NEP reactor 16 while retaining the aromatics for further processing and value confirmation. Figure 4 Shows Figure 1 An alternative embodiment of the embodiment of retaining aromatic hydrocarbons. 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 operation and configuration of the embodiment are basically the same as Figure 1 Same as above, with the following exceptions.

[0082] The NEP separation unit 20+ includes a debutanizer that separates a heavy stream including C4 and possibly C5 paraffins from the depropanizer bottoms stream in the debutanizer overhead line for recovery to the NEP reactor 16 in line 12+ and separates C5+ or C6+ aromatics in debutanizer bottoms line 28. The aromatics in the debutanizer bottoms line 28 may be further processed for valuable aromatics valuation.

[0083] The foregoing disclosure provides a process for converting naphtha into ethane and propane feeds wherein conversion to ethylene and propylene is maximized in ethylene and propylene production units, respectively.

[0084] Example

[0085] Example 1

[0086] Zeolite MFI with a Si / Al2 ratio of 40 is synthesized by crystallizing a mixture consisting of silica, alumina, alkali metals and OSDA containing tetrapropylammonium (TPA). The synthesized zeolite MFI is first mixed with AlPO4 sol and a gelling agent, and the resulting slurry is dripped into hot oil. The size of the spheres is controlled to obtain a calcined support of 1 / 16 inch diameter. The spheres coming out of the spheroidization section are separated from the oil and conveyed to the washing section for ammonia washing to remove sodium ions from the zeolite. The wet spheres are dried and then calcined to remove OSDA. Optionally, the calcined support is subjected to NH4 ion exchange using ammonium nitrate solution to remove residual sodium ions, followed by a drying step at 110°C before the metal incorporation step. Platinum incorporation is performed by an evaporative impregnation technique, by contacting the support with a solution of a platinum-containing compound (such as tetraamine nitrate platinate) in a rotary evaporator, followed by drying with heat introduced from a steam jacket. Once the spheres rolled freely, the impregnated support was steam oxidized in a quartz tube in a 3-zone furnace with 14.3 mol% steam in flowing air at 607°C for 2 hours. The platinum levels on the finished catalysts were 0.02 wt% and 0.40 wt% and were designated Examples 1.1 and 1.2, respectively.

[0087] Example 2

[0088] Using the alternative preparation sequence of Example 1.2, the oil drop spheres bound to AlPO4 were first calcined, subjected to ammonium ion exchange to remove residual sodium, and then steam calcined at 607°C for 2 hours using 14.3 mol% steam. The steam calcined support was impregnated with a tetraamine platinate solution by evaporation, followed by dry oxidation at 520°C for 2 hours in flowing air. The catalyst contained 0.4 wt% platinum and was designated Comparative Example 2.

[0089] Example 3

[0090] A series of catalysts containing 0 wt %, 0.005 wt %, 0.02 wt %, 0.1 wt % and 0.4 wt % platinum on AlPO4 bound MFI with a Si / Al2 ratio of 40 were prepared following the procedure described in Example 1, except that 7.7 wt % steam was co-fed during the steam oxidation step instead of 14.3 wt %. The catalysts were designated Examples 3.1, 3.2, 3.3, 3.4 and 3.5.

[0091] Example 4

[0092] 1 / 16 inch cylindrical extrudates were prepared from a formulation of 65% MFI-40, 20% Hisil 532SiO2, and 15% Ludox AS-40SiO2 prepared according to Example 1. An extrusion aid of 1% Methocel A was incorporated into the formulation. The extrudates were dried at 110°C and calcined at 520°C in flowing air in a muffle furnace. The calcined support was ammonium ion exchanged using a 1.0M ammonium nitrate solution to remove residual sodium. The ammonium ion exchange support was dried at 110°C and then tetraamine platinate was incorporated onto the support by an evaporative impregnation technique to obtain 0.2% or 0.4% platinum on the finished catalyst. The impregnated support was steam oxidized at 607°C and 14.3% steam for 2 hours and the final catalysts were designated Examples 4.1 and 4.2, respectively.

[0093] In a preparative displacement, 0.2 wt% Pt impregnated support was oxidized in dry flowing air at 520°C for 2 hours and designated Catalyst Comparative Example 4.3.

[0094] In another preparative permutation, a 0.2 wt% platinum impregnated support was prepared using chloroplatinic acid instead of tetraamine platinate nitrate, followed by steam oxidation at 607°C and 14.3% steam for 2 hours and designated Catalyst Example 4.4.

[0095] Example 5

[0096] In an example of preparing a zirconium oxide-bound zeolite catalyst, 366 g of zirconium oxyhydroxynitrate (ZHN) or ZrO(OH)(NO3) solution was mixed with 202 g of water and cooled to 10°C to 15°C. Subsequently, 60 g of a precooled gelling agent consisting of a 42% hexamethylenetetramine (HMT) solution was added under vigorous stirring at 10°C. The solution was then added to a ball mill. Thereafter, 229 g of H-MFI-Q powder was added to a ball mill containing the aforementioned solution, and the resulting slurry was ball milled for 30 minutes. The slurry was then dripped into 95°C oil to form gel spheres, which were kept in hot oil for two hours to age at 95°C. The aged spheres were then transferred to a washing tower and the oil was replaced with NH4NO3 at 95°C for more than 30 minutes. The spheres were then washed with NH4NO3 solution in the same tower, and the pH was adjusted with NH4OH at 95°C for 1 hour. The spheres were then dried overnight at 95°C. The spheres of the oil droplets were further dried at 110° C. for 2 hours and then calcined at 560° C. for 4 hours. The calcined supports were impregnated with solutions of chloroplatinic acid (CPA) or tetraamineplatinate nitrate (TAPN) by evaporation and subsequently oxidized at 520° C. in flowing dry air or at 607° C. in flowing air containing 14.3 mol % moisture, respectively. They are labeled as catalyst examples 5.1 (CPA dry oxide), 5.2 (CPA, steam oxide), 5.3 (TAPN, dry oxide) and 5.4 (TAPN, steam oxide).

[0097] Example 6

[0098] The MFI-40 material synthesized as described in Example 1 was mixed with silica sol, and the resulting mixture was dropped into a hot oil bath via a vibrating tip to form spheres. The spheres were aged, washed with water, dried and calcined in flowing air. Platinum was incorporated into the support via evaporative impregnation using a tetraamine nitrate platinate solution to obtain 0.02% by weight of platinum on the finished catalyst, followed by oxidation at 500°C in flowing air for 4 hours. The catalyst was designated as Catalyst Example 6.1. A related catalyst was prepared using a silicalite zeolite with a 500 Si / Al2 ratio. Platinum was incorporated into the support via evaporative impregnation using a tetraamine nitrate platinate solution to obtain 0.1% by weight of platinum on the finished catalyst, followed by granulation into 20×60 mesh and oxidation at 500°C in flowing air for 4 hours. The catalyst was designated as Catalyst Example 6.2.

[0099] Example 7

[0100] Molecular sieve AEL or SAPO-11 was obtained from Clariant. The composition, as measured by nitrogen adsorption, was 23.2 aluminum, 22.2 phosphorus, and 2.73 silicon, with a molecular weight of 256 m 2 / g BEA surface area, 0.11 ml / g micropore volume and 0.21 ml / g total pore volume. The molecular sieves were extruded with an alumina binder into a 70 / 30 SAPO-11 / Al2O3 formulation into 1 / 16 inch diameter cylinders followed by a final treatment with chloroplatinic acid to obtain 0.25 wt% platinum on the finished catalyst.

[0101] Example 8

[0102] Catalyst Example 1.1 was catalytically tested in a plug flow tubular reactor. The finished catalyst was crushed and 1 gram of 20×60 mesh catalyst was premixed with 3 grams of sand before loading into the reactor. The composition of the feed blend is shown in Table 1, simulating the composition of light naphtha. The feed was further spiked with dimethyl disulfide and thiophene to obtain a total sulfur content of 200 wppm. The catalyst was pre-reduced at 520°C for 3 hours in a flowing hydrogen stream and then reduced to the target reaction temperature before introducing the feed. At a temperature of 400°C to 430°C, a WHSV of 4.0 to 8.0, a molar ratio of hydrogen to hydrocarbon of about 1.0, and a total pressure of 200 psig, 13.8 bar (gauge pressure), it can be seen from Table 1 that, in terms of space velocity, the total amount of C2-C4 is always more than 80%, and the total amount of C2-C3 is always more than 70%, while keeping methane less than 5 weight %. In addition, of the C2-C4 produced, more than 50 wt% is C3. The amount of ethane can be increased via increased temperature while maintaining the yield of propane at 60 wt%. The conversion is calculated as the ratio of C3 hydrocarbons and aromatics in the product to C4+ hydrocarbons and non-aromatics in the feed.

[0103] Table 1

[0104]

[0105]

[0106] Example 9

[0107] The tests described in Example 8 were switched to a feed blend containing no sulfur. Under comparable test conditions, the catalysts exhibited comparable performance to the sulfur-containing feed blends, as shown in Table 2. For example, the performance results obtained after 13.6 hours of reaction using a sulfur-containing feed were comparable to the performance results obtained after 120.6 hours and 124.6 hours of reaction using a sulfur-free feed. Furthermore, increasing the temperature will allow fine-tuning of ethane production while maintaining propane yield at approximately 60%. It should be noted that increasing the contact time by decreasing the WHSV allows a propane yield of 65% to be achieved at lower temperatures while maintaining methane production below 3%.

[0108] Table 2

[0109]

[0110]

[0111] Example 10

[0112] The catalyst of Example 1.1 was tested following the same protocol described in Example 8. The testing was conducted by varying the molar ratio of hydrogen to hydrocarbon while keeping the pressure fixed at about 200 psig, 13.8 barg. As shown in Table 3, it was observed that 80 wt% C2 to C4 and greater than 70 wt% C2 to C3 were produced in each run. Moreover, ethane accounted for more than 60 wt% of the total amount of C2 to C3 and C2 to C4 because the molar ratio of hydrogen to hydrocarbon was reduced to below 5.0. In addition, by reducing the WHSV, the amount of ethane was further increased with increasing contact time. The increase in ethane production was also accompanied by a decrease in methane production and an increase in the production of valuable aromatics.

[0113] Table 3

[0114]

[0115]

[0116] Embodiment 11

[0117] The catalyst prepared according to Example 1.2 was tested following the same protocol as described in Example 8. The performance evaluation was performed by varying the total pressure, contact time expressed in WHSV and temperature from 480°C to 495°C while fixing the molar ratio of hydrogen to hydrocarbon at about 2.0. It is apparent that with reduced total pressure, greater than 60 wt% ethane and less than 10% methane can be obtained from C2 to C4 while maintaining 80% C2 to C4 by operating at a total pressure of 8 bar or less combined partial pressure of naphtha and H2. All pressures in the tables are absolute pressures.

[0118] Table 4

[0119]

[0120]

[0121] At equal pressure, higher temperature is favorable to ethane yield. At equal temperature, lower pressure is favorable to ethane yield.

[0122] Example 12

[0123] The catalysts of Examples 1.2 and 2 were tested following the same protocol as described in Example 8. Steam oxidation after platinum incorporation as in Catalyst Example 1.2 provided significantly higher ethane yields (greater than 50%) than Catalyst Example 2 prepared by steaming the support, followed by platinum incorporation, and then dry oxidation. Moreover, the total conversion of C2 to C3 produced by Catalyst Example 1.2 was greater than 80 wt%, with C2 accounting for more than 60 wt% of the C2 and C3 aggregates, and the methane yield was less than 8 wt% and less than 6 wt%. By comparison, Catalyst Example 2 resulted in lower ethane, lower total C2 to C3, and higher aromatics.

[0124] Table 5

[0125]

[0126]

[0127] Example 13

[0128] The catalysts of Examples 3.1 to 3.5 were tested in the temperature range of 390°C to 500°C at a hydrogen to hydrocarbon molar ratio of 2.0 according to a similar protocol as described in Example 8 using an apparatus consisting of multiple parallel reactors controlled at a pressure of 6 bar. 1 ml of the 20×60 mesh catalyst was charged into a separate reactor without diluent. The catalyst was pre-reduced at 520°C for 3 hours, then cooled to the first test temperature and pressurized to the test pressure. The results at 480°C are shown in Table 6.

[0129] Table 6

[0130] Example Feed 3.1 3.2 3.3 3.4 3.5 Platinum, weight % 0 0.005 0.02 0.1 0.4 WHSV, hr-1 2 2 2 2 2 H2 / HCBN molar ratio 2.0 2.0 2.0 2.0 2.0 Temperature, °C 477.8 478.1 479.3 478.3 480.6 Reaction time 88.6 89.0 87.8 89.4 88.2 Reactor outlet pressure, bar 5.6 5.6 5.9 5.7 5.7 C1 2.35 7.60 8.12 7.47 7.02 C2 6.35 46.62 56.80 65.76 65.08 C3 43.75 29.91 22.58 14.35 13.02 C2+C3 50.10 76.53 79.38 80.11 78.11 C4 17.45 2.98 2.48 1.94 1.72 C5 35.24 0.48 0.00 0.00 0.00 0.00 C2-C5 olefins 6.39 0.88 0.95 1.03 1.01 C6 47.57 12.63 0.15 0.12 0.06 0.05 Cyclohexane 15.64 C7 1.56 Aromatics 14.20 11.05 8.56 8.37 10.55 A9 1.48 0.46 0.22 0.22 0.29 A10 0.32 0.00 0.00 0.00 0.00 Heavy Aromatics 0.97 1.22 1.12 1.85 2.27 sum 100 100.00 100.00 100.00 100.00 100.00 <![CDATA[C2 / (C2-C4)%]]> 9.4 58.6 69 80 81.5

[0131] The results show that at a pressure of about 6 bar, a hydrogen to hydrocarbon molar ratio of 2.0, a platinum content as low as 50 wppm and 480°C, the presence of platinum is required to obtain a high C2 yield of more than 40 wt%. The presence of platinum is most critical to obtain a considerable C2 yield of more than 10 wt%. Increasing the platinum content by 0.1 wt% and increasing the higher C2 yield levels to more than 60 wt%, while maintaining a low level of methane formation of less than 8%. It should also be noted that by using a platinum-containing zeolite catalyst, while maintaining a combined C2 and C3 of more than 80 wt% and a methane formation rate of less than 8 wt%, a C2 / C3 ratio of more than 5 can be achieved.

[0132] Embodiment 14

[0133] The catalysts prepared according to Examples 4.1 to 4.4 were tested according to the test protocol described in Example 14, wherein the catalysts were tested in multiple parallel reactor devices using a feed simulating the light naphtha composition shown in Table 7 below using a catalyst loading of 1 ml of a 20×60 mesh size. The sixth column represents the cyclohexane and methylcyclopentane in the feed.

[0134] Table 7. Simulated Commercial Light Naphtha

[0135] Components Positive C5 Iso C5 Positive C6 Iso C6 CH&mCP Positive C7 sum weight% 23.09 11.54 26.40 21.10 15.82 2.05 100.00

[0136] Table 8

[0137] catalyst Binder Platinum Precursor Steam or dry oxidation Pt, wt% H2 / Pt C2 yield, % CH4,% Example 1.2 AlPO4 TAPN H2O 0.40 0.04 54.69 5.79 Comparative Example 2 AlPO4 TAPN dry 0.40 0.35 54.54 8.45 Example 4.1 SiO2 TAPN H2O 0.40 0.13 54.42 7.23 Example 4.2 SiO2 TAPN H2O 0.20 0.31 54.77 7.90 Comparative Example 4.3 SiO2 TAPN dry 0.20 0.54 54.90 13.00 Example 4.4 SiO2 CPA H2O 0.20 0.13 54.95 5.59

[0138] The test results are reported at comparable ethane production levels. SiO2-bound Catalyst Example 4.1 produced the highest methane formation rate among the four silica-bound catalysts. Catalyst Example 4.4, prepared using chloroplatinic acid as the platinum precursor in evaporative impregnation followed by steam oxidation at 14.4% steam and 607°C for 2 hours, gave the lowest methane production at 55% ethane yield. Catalyst Example 4.3, prepared using tetraammonium nitrate as the platinum precursor in evaporative impregnation followed by dry oxidation at 520°C for 2 hours, again produced significantly higher methane, as shown in Table 8.

[0139] Embodiment 15

[0140] The catalysts prepared according to Examples 5.1 to 5.4 were tested according to the testing protocol described in Examples 13 and 14, wherein the catalysts were tested in multiple parallel reactor devices using a feed simulating the light naphtha composition shown in Table 7 using a catalyst loading of 1 ml of a 20×60 mesh size.

[0141] Table 9

[0142] Example Catalyst 1.2 5.1 5.2 5.3 5.4 Pt Precursor TAPN CPA CPA TAPN TAPN Carrier AlPO4 ZrO2 ZrO2 ZrO2 ZrO2 Oxidation H2O dry H2O dry H2O <![CDATA[WHSV_hr -1 ]]> 2.0 2.0 2.0 2.0 2.0 H2 / LN mole 2.0 2.0 2.0 2.0 2.0 Temperature, C 450.0 450.0 450.0 450.0 450.0 Response time, hours 70.3 69.9 68.3 69.5 68.7 Outlet pressure, bar 5.2 5.3 5.7 6.4 5.9 Yield, % C1 4.71 4.96 4.54 5.83 4.13 C2 41.69 56.02 37.33 52.09 30.21 C3 41.32 26.53 47.96 26.36 51.82 C2+C3 83.01 82.55 85.29 78.46 82.03 C4 4.79 3.17 4.64 2.81 4.64 C5 0.00 0.00 0.00 0.00 0.00 C2-C5 olefins 0.75 0.48 0.66 0.50 0.54 BTX 6.11 6.57 4.30 9.54 7.73 A9 0.13 0.25 0.11 0.28 0.13 A10 0.00 0.00 0.00 0.00 0.00 Heavy Aromatics 0.51 2.02 0.47 2.59 0.80 sum 100.00 100.00 100.00 100.00 100.00 C2 / (C2-C4)% 47.49 65.36 41.51 64.10 34.86

[0143] Under relatively mild process conditions of 450°C, the ZrO2-bound catalysts produced very high ethane yields and very high ethane fractions in the C2 to C4 yields. As shown in Table 9, Catalyst Examples 5.1 and 5.3 produced 15 wt% higher ethane yields and greater than 60 wt% ethane in the combined total C2 to C4 compared to Catalyst Example 1.2. Dry oxidation after platinum incorporation consistently produced higher ethane yields and higher ethane fractions in the total C2 to C4 product compared to the AlPO4-bound Catalyst Example 1.2, regardless of whether chloroplatinic acid or tetraamineplatinate was used as the platinum precursor, than the counterparts prepared using steam oxidation after the platinum incorporation step.

[0144] Example 16

[0145] Catalyst Examples 6.1 and 6.2 were tested according to the protocol described in Example 9 using the feed compositions in Table 10.

[0146] Table 10

[0147] Components nC5 nC6 iC6 CH Methylcyclopentane nC7 weight% 38.44 43.56 0.66 13.42 0.25 2.06

[0148] Table 11

[0149]

[0150] As shown in Table 11, by operating at a higher WHSV with a shorter contact time using a low activity zeolite with a higher Si / Al2 ratio, such as in Catalyst Example 6.2, the process can obtain greater than 40% propane, with propane being predominantly in the C2-C3 aggregate, or having a high C3 / C2 (w / w) ratio of 3 to 10, as compared to Catalyst Example 6.1 which operates at a lower WHSV with an increased contact time.

[0151] Embodiment 17

[0152] Catalyst Example 7 comprising SAPO-11 was tested according to the protocol described in Example 8 using the feed composition in Table 10.

[0153] Table 12

[0154]

[0155] Catalyst Example 7 consisted of SAPO-11 and consistently exhibited greater than 10% aromatics, with benzene being the predominant aromatic over a wide range of temperatures and contact times. High benzene yields were characterized by the facile conversion of cyclohexane to benzene due to the elliptical pore shape of SAPO-11. It is contemplated that SAPO-11 could be employed in order to selectively convert cycloparaffinic components to aromatics, particularly in feeds containing high cycloparaffins.

[0156] Embodiment 18

[0157] Catalyst Example 1.2 was evaluated to show its utility in converting full range naphtha (FRN) to an effluent having a propane-rich, ethane-rich, or C2 to C4 stream having equal abundance of ethane and propane. As shown in Table 13, FRN was converted to an effluent containing greater than 40% propane at temperatures between 350°C and 460°C and a H2 / hydrocarbon molar ratio between 2.5 and 6.0. As shown in Table 14, FRN was converted to an effluent containing greater than 40% ethane at temperatures between 450°C and 500°C and a H2 hydrocarbon molar ratio between 1.0 and 3.0.

[0158] Table 13

[0159] Time, hours 37 42 50 56 60 70 77 90 96 108 Bed temperature, °C 349 375 375 375 401 401 426 451 451 451 <![CDATA[WHSV,hr -1 ]]> 1.93 1.93 1.93 1.92 1.92 1.92 1.92 1.92 1.91 1.96 H2 / HCBN 2.83 2.84 4.25 5.67 5.67 2.84 4.25 5.69 4.26 2.78 Reactor inlet pressure, psig 49.3 49.2 71.2 92.6 92.6 48.8 70.4 92.5 70.4 49.1 Feed C1 0.33 0.68 0.74 0.78 1.54 1.33 2.75 5.29 5.05 4.46 C2 3.44 5.95 5.78 5.72 9.37 9.88 14.66 21.32 22.25 26.26 C3 34.4 43.7 40.5 39.0 45.4 51.2 50.8 48.1 48.3 43.7 C4 23.2 20.1 22.3 23.2 19.4 13.3 10.0 5.9 4.6 3.8 C5 0.06 7.67 4.17 5.58 6.47 3.03 1.55 0.62 0.17 0.14 0.18 C6 10.82 3.09 0.89 1.12 1.29 0.26 0.21 0.08 0.05 0.06 0.07 mCP+CH+MCH 13.08 0.70 0.26 0.29 0.34 0.14 0.13 0.04 0.01 0.01 0.01 C7-C9 NA 59.9 13.1 10.1 10.0 10.1 7.3 8.5 7.4 6.3 6.4 7.5 1R aromatics 13.8 13.4 13.5 13.2 12.8 13.1 13.3 13.0 12.5 12.6 13.2 benzene 0.32 0.90 0.93 0.89 0.82 0.98 1.10 1.19 1.45 1.58 1.98 Toluene 2.79 4.01 4.03 3.92 3.83 4.04 4.27 4.37 4.49 4.65 5.03 Xylene 6.14 5.74 5.84 5.71 5.67 5.59 5.56 5.15 4.55 4.43 4.15 Ethylbenzene 0.46 0.03 0.03 0.03 0.03 0.01 0.01 0.01 0.00 0.00 0.00 A9+ 4.07 2.70 2.73 2.66 2.50 2.47 2.35 2.31 2.01 1.95 2.03 Heavier objects 2.33 0.63 0.65 0.52 0.38 0.45 0.63 0.68 0.41 0.48 0.83 sum 100 100 100 100 100 100 100 100 100 100 100 C2-C4 61.1 69.7 68.5 67.8 74.2 74.4 75.5 75.3 75.2 73.7 C2-C3 37.9 49.6 46.3 44.7 54.8 61.1 65.4 69.4 70.6 69.9 C2 / C3 0.10 0.14 0.14 0.15 0.21 0.19 0.29 0.44 0.46 0.60

[0160] Table 14

[0161] Time, hours 114 120 126 138 144 150 155 162 168 174 Bed temperature, °C 451 451 451 466 466 481 481 481 491 491 WHSV, hr-1 1.96 1.96 1.96 1.95 1.95 1.95 1.95 1.94 1.94 1.94 H2 / HCBN 2.79 2.08 1.38 2.09 2.80 2.79 2.07 1.39 1.41 2.11 Reactor inlet pressure, psig 74.2 74.0 73.9 74.0 74.0 74.1 74.0 74.7 74.6 74.5 Feed C1 5.63 5.05 2.96 5.14 6.21 8.34 6.63 4.36 5.04 6.55 C2 28.1 41.8 44.4 40.5 31.0 44.2 53.7 48.8 50.9 46.4 C3 42.1 27.6 18.5 27.4 38.4 23.6 13.5 11.3 11.5 19.0 C4 3.5 3.4 2.9 3.3 3.3 2.2 1.6 1.5 1.5 2.2 C5 0.06 0.16 0.29 0.32 0.27 0.16 0.11 0.12 0.13 0.13 0.15 C6 10.82 0.06 0.08 0.09 0.09 0.06 0.05 0.06 0.07 0.06 0.07 mCP+CH+MCH 13.08 0.01 0.02 0.04 0.02 0.01 0.01 0.01 0.01 0.01 0.01 C7-C9 NA 59.9 6.9 6.3 6.7 7.0 6.9 6.6 6.4 6.9 6.0 6.6 1R aromatics 13.8 12.8 14.1 19.1 14.9 12.8 13.7 16.4 22.0 20.2 16.9 benzene 0.32 2.04 2.63 3.75 2.76 2.12 2.90 3.82 4.47 4.68 3.45 Toluene 2.79 4.91 5.54 7.48 5.81 4.95 5.34 6.68 8.45 8.37 6.64 Xylene 6.14 3.95 3.96 5.10 4.17 3.85 3.43 3.92 5.72 4.89 4.37 Ethylbenzene 0.46 0.00 0.21 0.34 0.21 0.00 0.15 0.28 0.44 0.37 0.40 A9+ 4.07 1.92 1.78 2.46 1.97 1.88 1.90 1.71 2.94 1.91 2.03 Heavier objects 2.33 0.71 1.40 4.88 1.41 1.14 1.25 1.61 4.96 4.65 2.10 sum 100 100 100 100 100 100 100 100 100 100 100 C2-C4 73.7 72.7 65.8 71.2 72.7 69.9 68.7 61.6 63.9 67.6 C2-C3 70.2 69.4 62.9 67.8 69.4 67.8 67.1 60.1 62.4 65.5 C2 / C3 0.67 1.52 2.40 1.48 0.81 1.88 3.99 4.32 4.42 2.44

[0162] Embodiment 19

[0163] Catalyst Example 1.2 was evaluated to show its utility in converting a mixture of equal weight percentages of n-butane and isobutane to an effluent having a C2 to C3 stream rich in propane, rich in ethane, or equal abundance of ethane and propane. As shown in Table 15, at temperatures between 370°C and 470°C and H2 / hydrocarbon molar ratios between 1.0 and 3.0, butane was converted to an effluent containing greater than 40% propane. As shown in Table 16, at temperatures between 450°C and 500°C, more preferably between 460°C and 490°C, and H2 / hydrocarbon molar ratios between 1.0 and 2.0, more preferably between 1.0 and 1.6, butane was converted to an effluent containing greater than 40% ethane. Increasing contact time at reduced WHSV produces greater than 40% or more preferably greater than 50% ethane in the lower temperature range.

[0164] Table 15

[0165]

[0166]

[0167] Table 16

[0168]

[0169]

[0170] Embodiment 20

[0171] As shown in Tables 17 and 18, Catalyst Example 1.2 was evaluated to show its utility in producing ethane-rich and propane-rich light paraffins in a flexible manner, respectively. -1 When operating at a space velocity of 1.3, reducing the ratio of hydrogen to hydrocarbon will increase ethane and reduce methane at 450° C. and 465° C., respectively. By adjusting the ratio of hydrogen to hydrocarbon, light paraffin production can be switched to primarily ethane, a balance of ethane and propane, or primarily propane. For example, when operating at a given hydrogen to hydrocarbon ratio of 1.3 and a total pressure of 75 psig, the space velocity can be adjusted to switch light paraffin production from primarily ethane at a lower space velocity to primarily propane at a higher space velocity, as shown in Table 18, or vice versa.

[0172] Table 17

[0173]

[0174]

[0175] Table 18

[0176]

[0177] Embodiment 21

[0178] According to the test procedure in Example 19, a 50 / 50 mixture of isobutane and n-butane was reacted over catalyst Example 1.2 to produce ethane and propane. Figure 5 A plot depicting the reaction of a 50 / 50 mixture of isobutane and n-butane to produce ethane and propane as a function of ethylene intermediates. The total amount of ethylene and ethane, represented by the squares, increases with increasing ethylene concentration and plateaus at about 60 wt%, while propane, represented by the flat rectangles, decreases with increasing ethylene concentration and plateaus at about 20 wt% at ethylene concentrations between 0.15 wt% and 0.20 wt%. The presence of ethylene indicates that conversion to ethylene intermediates occurs in the reaction of butane to ethane and propane.

[0179] Specific implementation plan

[0180] 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.

[0181] A first embodiment of the present disclosure is a method for converting naphtha to primarily ethane, comprising contacting a naphtha stream with a zeolite catalyst comprising a precious metal and hydrogen to produce a light paraffin stream comprising ethane and propane; wherein the catalyst comprises at least 0.005 wt % of precious metals at a hydrogen to hydrocarbon ratio of less than 5. An embodiment of the present disclosure is one, any one, or all of the first embodiment of this paragraph to the previous embodiments of this paragraph, wherein the hydrogen to hydrocarbon ratio does not exceed 2. An embodiment of the present disclosure is one, any one, or all of the first embodiment of this paragraph to the previous embodiments of this paragraph, wherein the catalyst comprises less than 0.5 wt % of precious metals. An embodiment of the present disclosure is one, any one, or all of the first embodiment of this paragraph to the previous embodiments of this paragraph, wherein the catalyst comprises at least 0.02 wt % of precious metals. An embodiment of the present disclosure is one, any one, or all of the first embodiment of this paragraph to the previous embodiments of this paragraph, wherein the catalyst comprises at least 0.1 wt % of precious metals. One embodiment of the present disclosure is one, any one or all of the first embodiment of this paragraph to the previous embodiments of this paragraph, and also includes a reaction temperature of at least 425°C and a reaction pressure of less than 20 bar. One embodiment of the present disclosure is one, any one or all of the first embodiment of this paragraph to the previous embodiments of this paragraph, and also includes a reaction temperature of at least 440°C. One embodiment of the present disclosure is one, any one or all of the first embodiment of this paragraph to the previous embodiments of this paragraph, wherein the precious metal is platinum. One embodiment of the present disclosure is one, any one or all of the first embodiment of this paragraph to the previous embodiments of this paragraph, wherein the feed contains sulfur.

[0182] A second embodiment of the present disclosure is a method for converting naphtha mainly into propane, comprising contacting a naphtha stream with a zeolite catalyst comprising a noble metal and hydrogen to produce a light paraffin stream comprising ethane and propane; wherein the catalyst comprises at least 0.005% by weight of a noble metal, operating at a reaction temperature not exceeding 440°C. An embodiment of the present disclosure is one, any or all of the second embodiment of this paragraph to the previous embodiment of this paragraph, and also includes a noble metal catalyst concentration of 0.01% by weight to 0.1% by weight. An embodiment of the present disclosure is one, any or all of the second embodiment of this paragraph to the previous embodiment of this paragraph, wherein the ratio of hydrogen to hydrocarbon is less than 5. An embodiment of the present disclosure is one, any or all of the second embodiment of this paragraph to the previous embodiment of this paragraph, wherein the reaction temperature is not less than 350°C. An embodiment of the present disclosure is one, any or all of the second embodiment of this paragraph to the previous embodiment of this paragraph, and also includes a reaction pressure of no more than 20 bar. An embodiment of the present disclosure is one, any or all of the second embodiment of this paragraph to the previous embodiment of this paragraph, and also includes a reaction temperature of no more than 425°C. An embodiment of the present disclosure is one, any one or all of the second embodiment of this paragraph to the previous embodiment of this paragraph, wherein the precious metal is platinum. An embodiment of the present disclosure is one, any one or all of the second embodiment of this paragraph to the previous embodiment of this paragraph, wherein the feed contains sulfur.

[0183] A third embodiment of the present disclosure is a method for converting naphtha to ethane and propane, comprising contacting a naphtha stream with a zeolite catalyst comprising a precious metal and hydrogen to produce a light paraffin stream comprising ethane and propane; wherein the catalyst comprises at least 0.005 wt % of the precious metal; switching from a predominantly ethane mode at a hydrogen to hydrocarbon ratio of less than 2 to a predominantly propane mode at a hydrogen to hydrocarbon ratio of at least 2; or switching from a predominantly propane mode at a reaction temperature not exceeding 440° C. to a predominantly ethane mode at a reaction temperature exceeding 440° C. An embodiment of the present disclosure is one, any one, or all of the third embodiment of this paragraph to the prior embodiments of this paragraph, wherein the catalyst comprises 0.01 wt % to 0.2 wt % of the precious metal. One embodiment of the present disclosure is one, any or all of the third embodiment of this paragraph to the previous embodiment of this paragraph, further comprising when switching to the main propane mode, increasing the reaction pressure from less than 7 bar when in the main ethane mode to not less than 7 bar, or when switching to the main ethane mode, reducing the reaction pressure from not less than 7 bar when in the main propane mode to less than 7 bar. One embodiment of the present disclosure is one, any or all of the third embodiment of this paragraph to the previous embodiment of this paragraph, further comprising when switching from the main ethane mode to the main propane mode, increasing the ratio of hydrogen to hydrocarbon to at least 2.5. One embodiment of the present disclosure is one, any or all of the third embodiment of this paragraph to the previous embodiment of this paragraph, further comprising when switching from the main propane mode to the main ethane mode, unloading a catalyst having less than 0.1 wt % of precious metals and loading a catalyst having at least 0.1 wt % of precious metals, or when switching from the main ethane mode to the main propane mode, unloading a catalyst having at least 0.1 wt % of precious metals and loading a catalyst having less than 0.1 wt % of precious metals.

[0184] A fourth embodiment of the present disclosure is a method for converting naphtha into ethane and propane, comprising contacting a naphtha stream with a zeolite catalyst comprising a precious metal and hydrogen to produce a light paraffin stream comprising ethane and propane; wherein the catalyst comprises at least 0.1 wt. % of the precious metal; switching from a predominantly ethane mode at a WHSV of less than 4 hr-1 to a predominantly propane mode at a WHSV of at least 4 hr-1; or switching from a predominantly propane mode at a WHSV of at least 4 hr-1 to a predominantly ethane mode at a WHSV of less than 4 hr-1.

[0185] 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.

[0186] 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 process for converting naphtha into primarily ethane, comprising: contacting a naphtha stream with a zeolite catalyst comprising a noble metal and hydrogen to produce a light paraffin stream comprising ethane and propane; wherein the catalyst comprises at least 0.005 wt % of a noble metal at a hydrogen to hydrocarbon ratio of less than 5.

2. The method of claim 1, wherein the ratio of hydrogen to hydrocarbon does not exceed 2.

3. The process of claim 1 wherein the catalyst comprises less than 0.5 wt% of a noble metal.

4. The method of claim 1 further comprising a reaction temperature of at least 425°C and a reaction pressure of less than 20 bar.

5. The process according to claim 1, wherein the feed contains sulfur.

6. A process for converting naphtha into ethane and propane, comprising: contacting a naphtha stream with a zeolite catalyst comprising a noble metal and hydrogen to produce a light paraffin stream comprising ethane and propane; wherein the catalyst comprises at least 0.005% by weight of a noble metal; switching from a predominantly ethane mode at a hydrogen to hydrocarbon ratio of less than 2 to a predominantly propane mode at a hydrogen to hydrocarbon ratio of at least 2; or Switch from a predominantly propane mode at a reaction temperature not exceeding 440°C to a predominantly ethane mode at a reaction temperature exceeding 440°C.

7. The process of claim 6, wherein the catalyst comprises 0.01 wt% to 0.2 wt% of a noble metal.

8. The method of claim 6, further comprising increasing the hydrogen to hydrocarbon ratio to at least 2.5 when switching from a primarily ethane mode to a primarily propane mode.

9. The method of claim 6, further comprising unloading a catalyst having less than 0.1 wt % of precious metals and loading a catalyst having at least 0.1 wt % of precious metals when switching from a primarily propane mode to a primarily ethane mode, or unloading a catalyst having at least 0.1 wt % of precious metals and loading a catalyst having less than 0.1 wt % of precious metals when switching from a primarily ethane mode to a primarily propane mode.

10. A method for converting naphtha to ethane and propane comprising contacting a naphtha stream with a zeolite catalyst containing a noble metal and hydrogen to produce a light paraffin stream containing ethane and propane, wherein the catalyst contains at least 0.1% by weight of the noble metal; -1 The WHSV is switched to the main ethane mode for at least 4 hours. -1 or from a primary propane mode under WHSV of at least 4 hours -1 The WHSV switches to the main propane mode in less than 4 hours -1 The main ethane mode under WHSV.

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