Unit for preparing ethane and propane from naphtha with hydrogen slippage

Abstract

A process for producing ethylene from a hydrocarbon feed entails separating a hydrocarbon feed comprising ethane, hydrogen and / or methane and hydrocarbons having 3 or more carbons into an ethane-rich stream comprising hydrocarbons having (2) carbons or less and a heavy stream comprising hydrocarbons having (3) carbons or more. The ethane-rich stream is passed to a process for producing ethylene. The ethane-rich stream has a molar ratio of ethane to combined hydrogen and methane of at least 1.2 and less than or equal to (10).

Description

[0001] Specification

[0002] Priority Claim

[0003] This application claims the benefit and priority of U.S. Provisional Patent Application No. 63 / 425,427, filed on November 15, 2022, the entire disclosure of which is incorporated herein by reference. Technical Field

[0004] The present invention generally relates to methods and apparatuses for preparing ethylene from hydrocarbon feeds and for preparing ethylene from hydrocarbon feeds derived from naphtha, and more particularly to separating a hydrocarbon feed into an ethane-rich stream and a propane-rich stream, slipping hydrogen and methane products into the ethane stream to reduce separation costs in a naphtha-to-ethane-and-propane process, and sending the ethane stream containing hydrogen and methane to a process for preparing ethylene. Background Art

[0005] Due to the high value of ethylene and propylene compared to fuels, the industrial trend is to change refining capacity to manufacture increased petrochemicals. Steam cracking of naphtha is the industrial standard for preparing ethylene and propylene from naphtha, but the yield of ethylene plus propylene is low, less than 60 wt%, and typically less than 50 wt%, depending on the naphtha composition.

[0006] The naphtha-to-ethane-and-propane process (NEP) produces two main products, ethane and propane, which can be fed to downstream ethane steam crackers and propane dehydrogenation units to selectively produce ethylene and propylene, respectively, at a higher overall yield than a naphtha steam cracker.

[0007] The ethane product from the NEP process contains hydrogen and methane that are expensive to separate. This separation requires deep refrigeration or membranes and recycle compressors. Such methods are expensive and thus reduce the beneficial effect of the higher yields achievable by the NEP process when ethane and propane are fed to downstream ethane crackers and propane dehydrogenation units.

[0008] The present invention is provided to solve the above and other problems and to provide advantages and aspects not provided by this type of method and apparatus of the prior art. A complete discussion of the features and advantages of the present invention is deferred to the following detailed description with reference to the accompanying drawings. Summary of the Invention

[0009] One aspect of the present disclosure relates to a method for preparing ethylene from a hydrocarbon feed. The method includes separating the hydrocarbon feed comprising ethane, hydrogen and / or methane, and hydrocarbons having 3 or more carbons into an ethane-rich stream comprising hydrocarbons having 2 or fewer carbons and a heavy stream comprising hydrocarbons having 3 or more carbons. The ethane-rich stream is passed to a process for producing ethylene. The ethane-rich stream has a molar ratio of ethane to combined hydrogen and methane of at least 1.2 and less than or equal to 10.

[0010] Another aspect of the present invention relates to an apparatus for separating naphtha into ethylene and propylene. The apparatus includes a first fluid feed source for an ethane steam cracker. The first fluid feed has a molar ratio of ethane to combined hydrogen and methane in the fluid feed of at least 1.2 and less than or equal to 10. A control device controls the molar ratio.

[0011] Other features and advantages of the present invention will become apparent from the following description taken in conjunction with the accompanying drawings. BRIEF DESCRIPTION OF THE DRAWINGS

[0012] To understand the present invention, reference will now be made by way of example to the accompanying drawings, in which:

[0013] Figure 1 is a schematic diagram of the method and apparatus of the present disclosure;

[0014] Figure 2 is Figure 1 a schematic diagram of the method and apparatus of an alternative embodiment of

[0015] Figure 3 is Figure 1 a schematic diagram of the method and apparatus of an alternative embodiment of

[0016] Figure 4 is Figure 1 a schematic diagram of the method and apparatus of an additional alternative embodiment of

[0017] Figure 5 is Figure 1 a schematic diagram of the method and apparatus of a further alternative embodiment of

[0018] Figure 6 is a schematic diagram of the method and apparatus of the present disclosure, which shows a slip stream according to the present disclosure; and

[0019] Figure 7 is Figure 6 a schematic diagram of an alternative method and apparatus of the method and apparatus according to DETAILED DESCRIPTION

[0020] While the present invention admits of many different forms of embodiments, the preferred embodiments of the invention are shown in the drawings and will be described in detail herein. It is to be understood that the disclosure is to be regarded as an exemplification of the principles of the invention and is not intended to limit the broad aspects of the invention to the illustrated embodiments.

[0021] Referring 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 feed stream in line 11, which is heated and charged into the naphtha-to-ethane and propane (NEP) reactor 16 to contact the NEP catalyst. The NEP reactor block 16 can include a plurality of reactors in parallel or in series. The naphtha stream can include C4 to C12 hydrocarbons, preferably having a T10 between 0 °C and 10 °C and a T90 between 70 °C and 180 °C. The naphtha feed stream can include olefins, normal paraffins, isoparaffins, naphthenes, 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 naphtha weight hourly space velocity should be between 0.5 hr-1 and 10 hr-1, and preferably between 1 hr-1 and 4 hr-1. Some reactor designs include an inert diluent to increase the total weight hourly space velocity. However, all designs have a naphtha weight hourly space velocity between 1 h-1 and 4 h-1, where the naphtha weight hourly space velocity is defined as excluding the inert components. The molar ratio of hydrogen to hydrocarbon should be 0.5 to 5, and the total pressure should be 0.1 MPa (absolute) to 2 MPa (absolute). Under these conditions, the C2-C4 yield always exceeds 80 wt%, while the methane yield is less than 16 wt%, suitably less than 14 wt%, and generally less than 12 wt%, and preferably not exceeding 10 wt%. In another example, the molar ratio of hydrogen to hydrocarbon should generally not exceed 5, suitably not exceed 3, and preferably not exceed 2. Under these conditions, ethane can account for more than 60 wt% of the total C2 to C3 and of the total C2 to C4 produced in the NEP reactor 16. The low hydrogen-to-hydrocarbon ratio promotes the desired reaction kinetics initiated by dehydrogenation.

[0022] The NEP catalyst dehydrogenates naphtha molecules to their olefin analogs, interconverts the olefins to lighter olefins, and hydrogenates the lighter olefins to produce a light paraffin stream including ethane and propane. Interconversion can mean that olefins are oligomerized to higher carbon olefins, and then these higher carbon olefins are cracked to lower carbon olefins. This chemical mechanism avoids or minimizes the hydrocracking and pyrolysis reactions that produce methane. Methane is an undesirable by-product that represents an opportunity to produce valuable ethane and propane and consumes excessive hydrogen.

[0023] The NEP catalyst for converting naphtha into ethane and propane may comprise a molecular sieve, which includes a large-pore or medium-pore molecular sieve, which includes 10-membered rings 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.

[0024] 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 the chemical formula NanA1nSi96-nO92·16H2O (0 < n < 10). The ZSM-5 zeolite may include a silica to alumina molar ratio of 20 to 1000, 20 to 800, 20 to 600, 25 to 400, 25 to 200, or 25 to 80. The ZSM-5 zeolite may include a crystal size in the range of 10 nm to 1500 nm, 20 nm to 800 nm, 30 nm to 500 nm, 40 nm to 400 nm, or 50 nm to 300 nm.

[0025] The NEP catalyst may include a bound zeolite. The binder may include oxides of aluminum, silicon, zinc, titanium, zirconium and mixtures thereof. The binder may include phosphates in the binder or phosphates of the aforementioned oxide binder materials. Preferably, the binder is silica. The MFI zeolite may be loaded in a silica-containing binder or an alumina-containing binder (such as aluminum phosphate).

[0026] 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 support with a diameter of 1 / 888 inch to 1 / 32 inch. Alternatively, the zeolite may be mixed with a silica-containing binder and extruded into an extrudate with a diameter of 1 / 32 to 1 / 4 inch. The extrudate 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 support may be subjected to ammonium ion exchange using an ammonium nitrate solution to remove residual sodium ions and dried at 110 °C.

[0027] The NEP catalyst comprises a metal on the catalyst. The metal may comprise a transition metal. In another example, the metal may comprise platinum, palladium, iridium, rhenium, ruthenium, and mixtures thereof. The metal may be a noble metal. A modifier metal may also be included on the catalyst. The modifier metal may comprise tin, germanium, gallium, indium, thallium, zinc, silver, and mixtures thereof. The modifier metal should be more concentrated on the binder than on the zeolite. Each of from 0.01 wt% to 5 wt% of the transition metal and the modifier metal may be on the catalyst.

[0028] The metal may be incorporated into the binder by evaporation impregnation. A platinum solution (such as tetraamineplatinum nitrate or chloroplatinic acid) may be contacted with a binder spherical or extruded carrier that has been calcined and ion exchanged in a rotary evaporator, followed by drying and oxidation.

[0029] The NEP catalyst comprises a metal on a binder spherical or extruded carrier of the catalyst. Preferably, there is more metal on the binder than on the zeolite. 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. The zeolite and / or the entire NEP catalyst is steam oxidized to remove the metal from the zeolite. The steam oxidation is preferably carried out after the metal has been added to the catalyst. The dried impregnated spherical or extruded carrier may be steam oxidized in air for a sufficient time to provide the NEP catalyst. Steam oxidation in air at a temperature of from 500 °C to 650 °C and a steam of from 5 mol% to 30 mol% for 1 hour to 3 hours may be suitable.

[0030] The NEP catalysts must be reduced to activate them to catalyze the NEP reaction. For example, the catalyst may be reduced in flowing hydrogen at 500 °C to 550 °C for 3 hours before contacting the feed.

[0031] After paraffin conversion, the light paraffin stream is discharged from the NEP reactor 16 into the effluent line 18. The light paraffin stream may comprise at least 40 wt% of ethane or at least 40 wt% of propane or at least 70 wt% and preferably at least 80 wt% of 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 16 wt%, suitably less than 14 wt%, more suitably less than 12 wt%, preferably less than 10 wt%, more preferably less than 9 wt%, and most preferably less than 8 wt% of methane.

[0032] The presence of sulfur in the NEP reaction does not significantly affect the conversion. There may be at least 200 wppm of sulfur in the feed without a significant effect on the conversion. The NEP catalyst can handle up to 400 wppm and possibly 500 wppm of sulfur without a significant effect on the conversion.

[0033] 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 overhead line and a C2+ paraffin stream in the 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 recycle 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 overhead line 24 and a C3+ paraffin stream in the deethanizer bottom line. The C3+ paraffin stream can then be fed to a debutanizer to produce a propane stream in the debutanizer overhead line 26 and a heavy paraffin stream that can include C4+ hydrocarbons in the recycle line 12. It should be noted that the order of these towers can vary depending on the application. For example, in some cases, the deethanizer or debutanizer can be the first, and the demethanizer can be placed on either overhead stream. The NEP separation unit 20 can take other forms.

[0034] For example, the NEP separation unit 20 can omit the demethanizer, and the light paraffin stream in line 18 can be fed to a deethanizer that produces a C2 stream in the deethanizer overhead stream line. The C2 overhead stream can be separated in a hydrogen purification unit to recover the hydrogen stream in line 22, and the remaining ethane and methane from the hydrogen purification unit can include or supplement the ethane stream in line 24. The hydrogen purification unit can include a membrane unit, and the hydrogen recovered from the membrane unit can be further purified in an absorption tower and then recycled to the NEP reactor 16 in line 22. In an additional alternative (see Figure 2 ), the C2 stream from the deethanizer can be charged to an ethylene production unit 30 where ethane is converted to ethylene, but methane and hydrogen pass inertly to be recovered in a downstream ethylene recovery unit and sent via line 34 to the NEP reactor 16.

[0035] The ethane stream in line 24 can be fed into an ethylene production unit 30 where the ethane in the ethane stream is converted into ethylene. The ethylene production unit 30 is a steam cracking unit. The ethane stream in line 24 can be combined with unconverted ethane recovered from the steam cracking unit. The ethane stream in line 24 can be diluted with steam to achieve the desired partial pressure of ethane entering the ethane cracking furnace. The amount of steam mixed with ethane can preferably be from 0.15 kg steam / kg ethane to 0.5 kg steam / kg ethane. If an equal volume of a suitable inert substance is substituted, the amount of steam mixed with ethane can be less than the preferred amount. According to the principles of the present invention, hydrogen and methane are suitable inert substances for steam substitution. 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 fed into the ethane steam cracking unit in the 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 leaving the furnace of 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 can also include additional distillation towers, amine scrubbing towers, compressors, expanders, etc. to separate the cracked stream into cracked streams rich in individual light olefins, where the main component is the ethylene stream in line 32. Based on the ethane stream in line 24, the ethylene stream can include an ethylene yield of at least 75 wt%, preferably at least 80 wt%. Among the other components in the cracked stream leaving the ethane steam cracking, the ethylene production unit 30 can be hydrogen, methane, propylene, butene, and pyrolysis gas. Each of these components can be recovered and further processed.

[0036] The ethylene stream and the propylene stream in line 32 can be recovered or transported to a polymerization plant, a chemical plant, or exported. The butene stream can be recovered and used to produce plastics or other petrochemical products by methods such as polymerization, or exported. Based on the ethane stream in line 32, a product recovery of at least 50 wt%, typically at least 60 wt%, and suitably at least 70 wt% of valuable ethylene, propylene, and butene products can be achieved from the ethane steam cracking unit 30.

[0037] The ethylene production unit 30 can be an oxidative dehydrogenation (ODH) unit. The ethane stream in line 24 can be charged into the ODH unit. The ethane ODH process is an alternative technology to ethane steam cracking or ethane pyrolysis for converting ethane to ethylene. Ethane ODH involves contacting an ethane feed and an oxygen source in an ODH reaction zone in the presence of an ODH catalyst under conditions for oxidatively dehydrogenating at least a portion of the ethane 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. Mixed metal oxide catalysts have been found to be effective for the oxidative dehydrogenation reaction.

[0038] The ODH reactor can use a mixed metal oxide catalyst and operate at a temperature of about 300 °C to about 900 °C and produce more than 90% of 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. The oxidative dehydrogenation of ethane using a mixed metal oxide catalyst can 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, and a space velocity of 1000 cm3 / (gcat·hr) to 5000 cm3 / (gcat·hr), where the molar ratio of ethane to oxygen is 1.5:1 to 2:1, with sufficient inert diluent to achieve safe operating conditions. It is believed that the MoVNbTe oxide catalyst (and other related materials) having an M1 type structure is the best catalyst for ethane ODH.

[0039] The oxidative dehydrogenation of ethane 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 a gas hourly space velocity of 1000 h-1 to 5000 h-1. The Mg6MnO8 oxide catalyst (and other related materials) has been identified as a preferred material for the ethane ODH process, where the catalyst is also the oxygen source.

[0040] The propane stream in line 26 can be charged into the propylene production unit 40, where the propane in the propane stream is converted to propylene. The propylene production unit 40 can be a paraffin dehydrogenation (PDH) unit. A 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.

[0041] The dehydrogenation reaction can be carried out in a fluidized manner such that a gas, which may contain the reactant alkane with or without a fluidizing inert gas, is distributed to the reactor in a way 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 a reduction in the activity of the catalyst. The dehydrogenation catalyst must then be regenerated in a regenerator. The regenerator can combust the coke from the dehydrogenation catalyst and fuel gas to ensure sufficient enthalpy in the dehydrogenation reactor to promote the endothermic reaction.

[0042] The selected dehydrogenation catalyst should minimize the cracking reaction and favor the dehydrogenation reaction. Catalysts suitable for use herein include active metals that can be dispersed in a porous inorganic carrier material such as silica, alumina, kaolin, zirconia, or clay. Examples of catalysts include alumina or silica-alumina containing gallium, noble metals, and alkali or alkaline earth metals.

[0043] The catalyst support comprises a carrier material, a binder, and optionally a filler material to provide physical strength and integrity. The carrier material may include alumina or silica-alumina. Colloidal silica or alumina sol can be used as the binder. Alumina or silica-alumina typically contains alumina in the γ, θ, and / or δ phases. The nominal diameter of the catalyst support particles can be from 20 microns to 200 microns, and the average diameter is from 50 microns to 150 microns. Preferably, the surface area of the catalyst support is from 85 m2 / g to 140 m2 / g.

[0044] The fluidized dehydrogenation catalyst can include a dehydrogenation metal on a support. The dehydrogenation metal can be one or a combination of transition metals. Noble metals are preferred dehydrogenation metals, such as platinum or palladium. Gallium is an effective metal for paraffin dehydrogenation. The metal can be deposited on the catalyst support by impregnation or other suitable methods, or can be included in the carrier material or binder during catalyst preparation.

[0045] The acid functionality of the catalyst should be minimized to prevent cracking and favor dehydrogenation. Alkali and alkaline earth metals can also be included in the catalyst to attenuate the acidity of the catalyst. Rare earth metals can be included in the catalyst to control the activity of the catalyst. Metals can be incorporated into the dehydrogenation catalyst at a concentration of from 0.001 wt% to 10 wt%. In the case of noble metals, it is preferred to use from 10 parts per million by weight (ppm) to 600 ppm by weight of noble metal. More preferably, it is preferred to use from 10 ppm to 100 ppm by weight of noble metal. The preferred noble metal is platinum. Gallium should be present in the range of from 0.3 wt% to 3 wt%, preferably from 0.5 wt% to 2 wt%. Alkali and alkaline earth metals can be present in the range of from 0.05 wt% to 1 wt%.

[0046] The regenerated catalyst can be contacted with the propane stream in line 26, perhaps with a fluidizing gas, to lift the propane stream and the 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 separator. The propylene product gas can be quenched with a cooling fluid to prevent overreaction to form undesired 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 the light gases can be recycled to the regenerator as fuel gas for combustion to provide the enthalpy for the reaction.

[0047] The propylene production unit can also employ a catalytic moving bed reactor. The reactor section can include several parallel or series-connected radial flow reactors heated by chargers and interstage heaters. The propane stream, possibly with added hydrogen, flows through the screened central tube through the annular dehydrogenation catalyst bed to the outer effluent ring in each dehydrogenation reactor. The flow can be in the opposite direction. The dehydrogenation catalyst can include noble metals and their mixtures, modifiers selected from the group consisting of alkali metals or alkaline earth metals and their compositions, components selected from the group consisting of tin, germanium, lead, indium, gallium, thallium and their compositions, and a porous support forming the catalyst particles. The catalyst support can include oil-drop alumina spheres.

[0048] The dehydrogenation conditions can include a temperature of 400 °C to 900 °C, a pressure of 0.01 absolute atmosphere to 10 absolute atmospheres, and a liquid hourly space velocity (LHSV) of 0.1 hr-1 to 100 hr-1. The pressure in the dehydrogenation reactor is maintained at the lowest feasible level consistent with equipment limitations to maximize the chemical equilibrium advantage. The spent dehydrogenation catalyst in the annular catalyst bed can be removed from the bottom of the bed and transferred to the regenerator to burn the coke from the catalyst in air at 450 °C to 600 °C. The noble metals on the catalyst can be redispersed, dried by an oxychlorination method, and returned to the top of the dehydrogenation catalyst bed as the regenerated dehydrogenation catalyst.

[0049] The dehydrogenation effluent from the propylene production unit 40 can be cooled, compressed, dried, and hydrogen is cryogenically separated from the hydrocarbons, with a net gas purity of hydrogen of 85 mol% to 93 mol%. The hydrocarbon liquid is selectively hydrogenated to convert dienes and acetylene, and the hydrocarbon liquid is 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. Propane can be recycled as feed to the propylene production unit 40.

[0050] The heavy stream that can be withdrawn from the bottom of the depropanizer in line 12 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-rich gas stream in line 22 and charged into the NEP reactor in line 11 to produce more ethane and propane.

[0051] According to an alternative, the light paraffin stream in line 18 can 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 can include isobutane. Here, the propane and isobutane stream in line 26 can be fed to the propylene production unit 40. The propylene production unit 40 can be equipped to dehydrogenate propane in the propane and isobutane stream 26 into propylene and dehydrogenate isobutane in the propane and isobutane stream in line 26 into isobutene in the same dehydrogenation reactor. The fractionation section from the propylene production unit 40 can include a depropanizer downstream of the bottom of the propylene-propane separation column to separate unreacted propane at the top from C4 hydrocarbons. An isobutene-isobutane separation column in communication with the downstream of the bottom of the depropanizer can provide isobutene at the top that can be recovered as a product, and unreacted isobutane at the bottom of the separator can be returned to the propylene production unit 40 to be converted into isobutene.

[0052] Unreacted normal C4+ hydrocarbons including normal C4 and higher paraffins in the heavy stream 12 can be recovered to be combined with the naphtha stream in the feed line 10 and the hydrogen stream in line 22 and charged into the NEP reactor 16 in line 11.

[0053] Figure 3 is shown Figure 1 an alternative of which employs a normal butene conversion unit 50. Elements in Figure 1 with the same configuration as in Figure 3 will have the same reference numerals as in Figure 1 . Elements in Figure 1 with a different configuration from the corresponding elements in Figure 3 have the same reference numerals but are denoted with an apostrophe ('). Figure 3 The configuration and operation of the apparatus of Figure 1 are substantially the same as in

[0054] except as follows. Figure 1 In addition to the hydrogen-rich stream in line 22, the ethane stream in line 24, and the propane and isobutane stream in line 26, the NEP separation unit 20' additionally provides a normal butane product stream in line 46. The normal butane stream in line 46 can be charged into the normal butene production unit 50 to convert normal butane in the normal butane stream into normal butene. This may be useful in cases where the propylene production unit cannot process propane and normal butane in the same dehydrogenation reactor. The normal butene production unit can be a dehydrogenation unit as described for the

[0055] The 1-butene production unit can feed a 1-butene stream into the 1-butene-1-butane separator column. The 1-butene stream in the overhead line of the separator can be regarded as the 1-butene product in line 52, while the unreacted 1-butane in the bottom line of the column can be recycled back to the 1-butene production unit 50 for conversion to 1-butene.

[0056] In some cases, the butene production unit 50 can be a butene production unit capable of converting both n-butane and isobutane into 1-butene and isobutene. In this case, line 46 carriers a propane-lean butane stream to the butene production unit 50, and the propane stream in line 26 carriers a butane-lean 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 butenes.

[0057] The heavy stream from the NEP separation unit 20' in line 12 may include a C5+ hydrocarbon stream, but it may include C4 hydrocarbons and include a C4+ hydrocarbon stream. The unreacted C4+ or C5+ hydrocarbons including C4 or C5 and heavier hydrocarbons in the heavy stream 12 can be recovered and combined with the naphtha stream in the feed line 10 and the hydrogen stream in line 22 and charged into the NEP reactor 16 in line 11.

[0058] The heavy paraffin stream in line 12 may also include aromatics, benzene, toluene, and xylene. Figure 4 is shown Figure 1 an alternative that maximizes the production of ethane and propane by hydrotreating the entire heavy stream in line 12” in a hydrotreating reactor 60 to saturate the aromatic rings to cycloalkanes and providing a recycle feed to the NEP reactor 16 in line 62. Elements having the same configuration as in Figure 1 will have the same reference numerals as in Figure 4 Elements having a different configuration from the corresponding elements in Figure 1 will have the same reference numerals but will be denoted with double primes (”). Figure 1 will have the same reference numerals but will be denoted with double primes (”). Figure 4 The configuration and operation of the equipment of Figure 4 are substantially the same as in Figure 1 except as follows.

[0059] Mix, heat, and charge the heavy stream including C4+ paraffins and aromatics, benzene, toluene, and xylene in line 12” with the hydrogen stream in line 64 into the hydrotreating reactor 60.

[0060] The hydrotreating reactor 60 may have one or more hydrotreating catalyst beds to saturate the aromatic rings in the heavy stream. The heavy stream can be charged to the hydrotreating reactor 60, and the hydrotreating inlet temperature can be in the range of 200 °C (392 °F) to 400 °C (752 °F). If more than one catalyst bed is used, the hydrotreating reactor 60 can employ an interbed hydrogen quench stream.

[0061] Suitable hydrotreating catalysts are any known conventional hydrotreating catalysts and include those hydrotreating catalysts composed 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 hydrotreating catalysts include zeolite catalysts and noble metal catalysts, where the noble metal is selected from palladium and platinum. Within the scope of this specification, more than one type of hydrotreating catalyst is used in the same hydrotreating reactor 60. The Group VIII metal is typically present in an amount in the range of 2 wt% to 20 wt%, preferably 4 wt% to 12 wt%. The Group VI metal will typically be present in an amount in the range of 1 wt% to 25 wt%, preferably 2 wt% to 25 wt%. Generally, the hydrotreating conditions include a pressure of 700 kPa (100 psig) to 21 MPa (3000 psig). The hydrotreating outlet temperature can be in the range between 300 °C (572 °F) and 427 °C (800 °F).

[0062] The saturated aromatics and C4+ paraffins in the hydrotreated heavy stream in line 62 can be recovered and 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. The hydrogen from the hydrotreating reactor 60 can be recovered with the saturated aromatics to the NEP reactor 16 to reduce or eliminate the hydrogen requirement from line 22.

[0063] The NEP separation unit 20+ may include a debutanizer tower that separates C4 and possibly C5 hydrocarbons for recovery to the NEP reactor 16 in line 12+, while retaining the aromatics for further processing and value confirmation. Figure 5 is shown Figure 1 an alternative for retaining aromatics. Having the same configuration as Figure 1 in Figure 5 the elements in Figure 1 will have the same reference numerals as Figure 1 in Figure 5 The elements in

[0064] The NEP separation unit 20+ includes a debutanizer that separates a heavy stream including C4 and possibly C5 paraffins from the debutanizer bottoms stream in the debutanizer overhead line for recycle to the NEP reactor 16 in line 12+ and separates C5+ or C6+ aromatics in the debutanizer bottoms line 28. The aromatics in the debutanizer bottoms line 28 can be further processed for value determination of valuable aromatics.

[0065] The present disclosure also relates to the concept of desiring to slip a fraction of hydrogen and methane into the ethane product to reduce separation costs in the naphtha-to-ethane and propane process (NEP). Slipping a fraction of hydrogen and methane into the ethane product reduces the total compression power (i.e., refrigeration power) of the NEP process and may require less equipment compared to alternatives.

[0066] Slipping hydrogen and methane into the ethane product also allows for more efficient heat recovery, which further reduces energy costs compared to alternatives. The principles of the present disclosure have the greatest beneficial effect when the hydrogen and methane in the ethane product have a minimal impact on the downstream process. In the case where the downstream process is ethane steam cracking, the ethane feed is diluted with steam to reduce the ethane partial pressure in the cracking furnace. According to the principles of the present disclosure, the hydrogen and methane slipped in the NEP process ethane product can partially replace the steam in the ethane steam cracker and have a minimal impact on ethylene yield as long as the volume of hydrogen plus methane is less than or equal to the total volume of steam diluent that would be used for a pure ethane feed.

[0067] The typical steam dilution rate in an ethane steam cracker is 0.3 kg steam / kg ethane. This corresponds to 33 volume % steam to 67 volume % ethane. According to the principles of the present disclosure, slipping hydrogen and methane in the NEP process ethane product is beneficial for an ethane to combined hydrogen plus methane volume ratio of at least 2 before considering ethane recovery.

[0068] The ethane steam cracker effluent typically contains unreacted ethane, which is recycled back to the ethane cracking reactor. This ethane is further diluted with steam at a typical rate of 0.3 kg steam / kg ethane recycle. According to the principles of the present disclosure, the hydrogen and methane slipped in the NEP process ethane product can also displace a portion of the steam required for ethane recovery. The steam required for ethane recovery can be up to two-thirds of the steam required for fresh ethane feed. After considering the steam required for ethane recovery, slipping hydrogen and methane in the NEP process ethane product is beneficial for an ethane to combined hydrogen plus methane volume ratio of at least 1.2.

[0069] In some cases, it may be advantageous to replace a portion of the steam that is less than the total desired amount. In these cases, for an ethane to combined hydrogen plus methane volume ratio of at least 1.5, it may be advantageous to slip hydrogen and methane in the NEP process ethane product.

[0070] The main alternatives to the principles of the present disclosure are to use a membrane and recycle compressor to separate hydrogen from the ethane product, use a demethanizer column to separate hydrogen and methane from the ethane product, use a cooling and flash configuration to separate hydrogen and methane from the ethane product, or use a dividing wall column deethanizer to separate hydrogen and methane from the ethane product, where ethane is a byproduct.

[0071] Membranes have been used to selectively remove hydrogen from hydrocarbon streams, but the hydrogen product is produced at low pressure and must be compressed before it can be used for any application other than as fuel. Employing a recycle compressor increases the compression power and capital cost, which can be avoided by implementing the principles of the present disclosure.

[0072] Alternatively, hydrogen and methane can be separated from the ethane product by cooling and flashing the ethane in several stages and / or separating hydrogen and methane in a demethanizer column. This requires a refrigerant at -80 °C or lower to achieve good separation. This increases the compression power of the refrigeration compressor and the capital cost of the cryogenic heat exchanger.

[0073] Dividing wall columns have also been used to achieve ternary separations, with a capital cost lower than that achievable by separating in a stepwise manner. In this case, a dividing wall distillation column separates propane, ethane, and a hydrogen / methane product. Separating hydrogen / methane from ethane requires a refrigerant temperature in the condenser lower than that required for a simple deethanizer column separating C2- from propane. Additional refrigeration compression power is required in this scenario compared to slipping hydrogen and methane in the ethane product.

[0074] Current standard practice is to separate ethane from hydrogen and methane at high purity, typically less than 5 wt% methane and no hydrogen. This corresponds to an ethane to combined hydrogen plus methane volume ratio greater than 10. Depending on the molar ratio of hydrogen, methane, and ethane in the feed, current standard practice can be achieved by the main alternatives to the principles of the present disclosure described above.

[0075] The principle of the present disclosure of slipping hydrogen and methane into the ethane product has the greatest beneficial effect when this reduces the separation cost in the NEP process. This advantage can occur when ethane is purified to a lower purity than standard practice. The reduction in separation cost can be achieved when the ethane to combined hydrogen plus methane volume ratio is less than or equal to 10, more preferably when the volume ratio is less than or equal to 8, and most preferably when the volume ratio is less than or equal to 5.

[0076] SeeFigure 6 and Figure 7 , in a method and apparatus for producing ethylene and propylene from a hydrocarbon feedstock, a naphtha stream is contacted with a catalyst in a NEP reactor 16 to transfer a light paraffin effluent in line 18 to a cold box 46. A refrigeration unit 50 transfers a refrigeration fluid in line 52 to maintain a desired temperature in the cold box 46 at -20°C. The cold box 46 can be a multi-stream heat exchanger.

[0077] The effluent stream is a hydrocarbon stream that contains hydrogen and C1 to C6+ hydrocarbons, preferably contains ethane, hydrogen, and / or methane and hydrocarbons having 3 or more carbons. It is cooled in the cold box 46 and transferred as a feed via line 18 to a NEP separator 20, such as a deethanizer column. The hydrocarbon stream is separated into a bottoms product that contains propane and other heavy hydrocarbons having 3 or more carbons. It can be transferred directly in line 26 to a propylene production unit 40, or further separated into a heavy hydrocarbon stream containing C4+ hydrocarbons and a propane-rich stream that is transferred in line 26 to the propylene production unit 40.

[0078] The NEP separator 20 also produces an overhead product that contains hydrogen and hydrocarbons having 2 or fewer carbons, preferably an ethane-rich stream containing hydrocarbons having 2 or fewer carbons. The overhead product is transferred and passed through the cold box 46 in line 56 to condense a portion of the overhead product. The cold box 46 serves as a condenser for the overhead product in line 56, and block 60 is a condenser receiver (also known as a reflux drum). Then, the two-phase overhead product passes through the condenser receiver 60. The liquid stream exits the condenser receiver 60 and returns to the NEP separator 20 via line 64.

[0079] An ethane-rich stream containing hydrogen and hydrocarbons having 2 or fewer carbons (preferably ethane, hydrogen, and methane) leaves the condenser receiver 60 via line 24. The ethane stream can be fed downstream via line 24 to a process for producing ethylene from the ethane-rich stream, such as an ethylene production unit 30, for example a downstream ethane steam cracker or an ethane oxidative dehydrogenation process. The molar ratio of ethane to hydrogen plus methane (ethane:(hydrogen + methane)) is important in such an ethane-rich stream of the present invention. The molar ratio should be at least 1.2, preferably at least 1.5, more preferably at least 1.8, and most preferably 2.0. At a molar ratio of at least 2.0, hydrogen and methane replace all steam diluents used for ethane feed before considering ethane recovery. When the molar ratio is greater than 2.0, greater beneficial effects are achieved because some steam diluents can be fed to the ethane cracker without increasing the size of the ethane cracker. At a molar ratio of 2.4, hydrogen and methane replace two-thirds of the steam diluents used for ethane feed before considering ethane recovery. Some steam is beneficial for the ethane steam cracker to prevent coke accumulation. The ethane steam cracker effluent contains unreacted ethane, which can be recovered and returned to the ethane steam cracker. This recovered ethane is typically also diluted with steam to maintain a preferred ethane partial pressure. This additional dilution steam can also be replaced by hydrogen and methane in the ethane-rich stream in line 24. At a molar ratio of ethane to combined hydrogen and methane of 1.2, hydrogen and methane replace all steam diluents for both fresh feed and ethane recovery. When the molar ratio is greater than 1.2, greater beneficial effects are achieved because some steam diluents can be fed to the ethane cracker without increasing the size of the ethane cracker.

[0080] A pressure control device or pressure reduction device 66 is located within line 24 downstream of the condenser receiver 60 and upstream of the multi-stream heat exchanger cold box 46. One purpose of the pressure reduction device 66 is to enhance heat recovery by first reducing the pressure of stream 24, which reduces the temperature of stream 24 as the gas expands, and then recovering additional heat in a heat exchanger. Thus, the ethane-rich stream overhead product is passed through the pressure reduction device 66 and then through a heat recovery device, such as a heat exchanger, for example the multi-stream heat exchanger cold box 46. This method recovers heat, which provides some of the cooling load to condense a portion of the deethanizer overhead distillate in line 56.

[0081] The pressure reduction device 66 can be a valve (see Figure 6 ) or a turbine, such as a turboexpander (see Figure 7 ). The pressure reduction device 66 is placed before the ethane-rich stream heat recovery exchanger so that more heat can be extracted from the ethane-rich stream, which reduces the net refrigeration requirement. Combining the hydrogen and methane slipstream into the ethane-rich stream can be combined with the pressure reduction device 66 to reduce utilities.

[0082] As described above, the pressure reducing device 66 can be a recovery turbine, such as a turboexpander. A turboexpander is a centrifugal or axial flow turbine through which high-pressure gas expands to produce work that is typically used to drive a compressor or generator. Here, the turboexpander can partially liquefy the ethane-rich stream. Compared with the case where the pressure reducing device 66 is a valve, the work recovered from the turboexpander in a heat recovery device (such as the multi-stream heat exchanger cold box 46) plus the heat recovered from the partially liquefied ethane-rich stream in line 24 can reduce the utility more. One purpose of the valve / turbine in the example is to enhance heat recovery by first reducing the pressure and then recovering additional heat in the heat exchanger.

[0083] Compared with the alternative, the proposed solution saves energy and some equipment while not increasing the size of the downstream ethane steam cracker of the ethylene production unit 30. Table 1 shows Figure 6 the pressure reducing valve, Figure 7 the comparison of savings between the pressure reducing turbine and the alternative solution.

[0084] Table 1: Energy Savings by the Method

[0085]

[0086] The term "in communication with" means that fluid flow is operably permitted between the listed components, which can be characterized as "in fluid communication".

[0087] The term "downstream" means that at least a portion of the fluid flowing towards the main body in a downstream communication can flow operably from an object in fluid communication with it.

[0088] The term "upstream" means that at least a portion of the fluid flowing out of the main body in an upstream communication can flow operably towards an object in fluid communication with it.

[0089] The term "in direct communication with" means that the fluid flow from an upstream component enters a downstream component without passing through any other intervening container.

[0090] The term "in indirect communication with" means that the fluid flow from an upstream component enters a downstream component after passing through an intervening container.

[0091] The term "bypass" means that the object loses downstream communication with the bypassed main body at least within the scope of the bypass.

[0092] The term "reflux" means the liquid that condenses from the overhead vapor of a distillation column and returns to the top of the column.

[0093] As used herein, the term "major" or "substantially" means greater than 50%, suitably greater than 75%, and preferably greater than 90%.

[0094] As used herein, the term "rich in" is defined as at least 50 mol%.

[0095] Although specific embodiments have been illustrated and described, many modifications can be envisioned without materially departing from the essence of the present invention, and the scope of protection is limited only by the scope of the appended claims.

[0096] Specific Embodiments

[0097] While the following is described in conjunction with specific embodiments, it should be understood that the description is intended to illustrate and not limit the scope of the foregoing description and the appended claims.

[0098] A first embodiment of the present invention is a method for preparing ethylene from a hydrocarbon feed, the method comprising separating a hydrocarbon feed comprising ethane, hydrogen and / or methane and hydrocarbons having 3 or more carbons into a lean ethane stream comprising hydrocarbons having 2 or fewer carbons and a heavy stream comprising hydrocarbons having 3 or more carbons; and passing the lean ethane stream to a process for producing ethylene, wherein the lean ethane stream has an ethane to combined hydrogen and methane molar ratio of at least 1.2 and less than or equal to 10. One embodiment of the present invention is one, any or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, the method further comprising separating the hydrocarbon feed in a deethanizer upstream of the process for producing ethylene, wherein the deethanizer outputs a propane stream as a bottoms product and a tops product, wherein the tops product comprises ethane, hydrogen and methane. One embodiment of the present invention is one, any or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, the method further comprising passing the tops product through a condenser receiver, and returning the liquid as reflux to the deethanizer, and passing the lean ethane stream to the process for producing ethylene. One embodiment of the present invention is one, any or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the lean ethane stream is passed through a pressure reducing device and then through a heat recovery device. One embodiment of the present invention is one, any or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the heat recovery device is a heat exchanger that provides part of the cooling load to condense part of the tops product of the deethanizer. One embodiment of the present invention is one, any or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the pressure reducing device is a valve. One embodiment of the present invention is one, any or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the pressure reducing device is a turbine. One embodiment of the present invention is one, any or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the turbine is a turboexpander that partially liquefies the lean ethane stream. One embodiment of the present invention is one, any or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the hydrocarbon feed is derived from naphtha in an upstream process. One embodiment of the present invention is one, any or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the molar ratio is at least 1.5. One embodiment of the present invention is one, any or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the molar ratio is at least 1.8.One embodiment of the present invention is one, any or all of the first embodiment to the foregoing embodiments of this paragraph, wherein the molar ratio is at least 2.0. One embodiment of the present invention is one, any or all of the first embodiment to the foregoing embodiments of this paragraph, wherein the molar ratio is less than or equal to 8. One embodiment of the present invention is one, any or all of the first embodiment to the foregoing embodiments of this paragraph, wherein the molar ratio is less than or equal to 5.

[0099] A second embodiment of the present invention is an apparatus for separating naphtha into ethylene and propylene, the apparatus comprising a first fluid feed source of an ethane steam cracker, the first fluid feed being an ethane-rich stream, the ethane-rich stream having a molar ratio of ethane to combined hydrogen and methane in the fluid feed of at least 1.2 and less than or equal to 10; and a pressure reduction device, wherein the pressure reduction device first reduces the pressure of the ethane-rich stream, which reduces the temperature of the ethane-rich stream as the gas of the ethane-rich stream expands, and then recovers heat in a heat exchanger. An embodiment of the present invention is one, any, or all of the second embodiment of this paragraph to the foregoing embodiments of this paragraph, the apparatus further comprising a deethanizer tower, wherein the deethanizer tower outputs a top product to the heat exchanger, and wherein the top product comprises ethane, hydrogen, and methane, and wherein optionally a propane-rich product is sent to a propane dehydrogenation unit to produce propylene. An embodiment of the present invention is one, any, or all of the second embodiment of this paragraph to the foregoing embodiments of this paragraph, the apparatus further comprising a condenser receiver, the condenser receiver receiving the top product from the heat exchanger and transferring the ethane-rich stream to the pressure reduction device, wherein the pressure reduction device receives the ethane-rich stream from the condenser. The apparatus according to claim 16, wherein the pressure reduction device is one of a valve or a turbine, and the first fluid feed is transferred through the pressure reduction device to the heat exchanger before the ethane steam cracker. An embodiment of the present invention is one, any, or all of the second embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the molar ratio is at least 1.5. An embodiment of the present invention is one, any, or all of the second embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the molar ratio is at least 1.8. An embodiment of the present invention is one, any, or all of the second embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the molar ratio is at least 2.0. An embodiment of the present invention is one, any, or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the molar ratio is less than or equal to 8. An embodiment of the present invention is one, any, or all of the first embodiment of this paragraph to the foregoing embodiments of this paragraph, wherein the molar ratio is less than or equal to 5.

[0100] Although no further elaboration is provided, it is believed that those skilled in the art can make the most of the present invention by using the foregoing description and can easily determine the basic features of the present invention without departing from the essence and scope of the present invention to make various changes and modifications thereto and adapt it to various uses and conditions. Therefore, the foregoing preferred specific embodiments should be understood as being merely illustrative and not in any way limiting the remainder of the disclosure, and are intended to cover various modifications and equivalent arrangements included within the scope of the appended claims.

[0101] In the foregoing, all temperatures are shown in degrees Celsius and all parts and percentages are by weight unless otherwise indicated.

Claims

1. A method for preparing ethylene from a hydrocarbon feedstock, the method comprising: separating the hydrocarbon feedstock comprising ethane, hydrogen and / or methane and hydrocarbons having 3 or more carbon atoms into a lean ethane stream comprising hydrocarbons having 2 or fewer carbon atoms and a heavy stream comprising hydrocarbons having 3 or more carbon atoms; and passing the lean ethane stream to a process for producing ethylene, wherein the lean ethane stream has a molar ratio of ethane to combined hydrogen and methane of at least 1.2 and less than or equal to 10.

2. The method according to claim 1, the method further comprising separating the hydrocarbon feedstock in a deethanizer upstream of the process for producing ethylene, wherein the deethanizer outputs a propane stream as a bottoms product and a tops product, and wherein the tops product comprises ethane and at least one of hydrogen or methane.

3. The method according to claim 2, the method further comprising passing the tops product through a condenser receiver, and returning the liquid as reflux to the deethanizer, and passing the lean ethane stream to the process for producing ethylene.

4. The method according to any one of the preceding claims, wherein the lean ethane stream is passed through a pressure reduction device and then through a heat recovery device.

5. The method according to claim 4, wherein the heat recovery device is a heat exchanger that provides a portion of the cooling load to condense a portion of the tops product of the deethanizer.

6. The method according to claim 5, wherein the pressure reduction device is selected from a valve, a turbine, and a turboexpander that partially liquefies the lean ethane stream.

7. The method according to any one of the preceding claims, wherein the hydrocarbon feedstock is derived from naphtha in an upstream process.

8. The method according to any one of the preceding claims, wherein the molar ratio is at least 1.

5.

9. The method according to any one of the preceding claims, wherein the molar ratio is at least 1.

8.

10. The method according to any one of the preceding claims, wherein the molar ratio is less than 8.

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