NL9301130A - INTEGRATED MEMBRANE / HYDROCRAFT PROCESS FOR IMPROVED USE OF SUPPLY MATERIAL IN THE PREPARATION OF LOW EMISSION GASOLINE. - Google Patents

Description

INTEGRATED MEMBRANE / HYDROCRAQUE PROCESS FOR IMPROVED USE OF SUPPLY MATERIAL IN THE PREPARATION OF LOW-EMISSION GASOLINE

Background of the Invention

Field of the Invention

The present invention relates to a process for an improved utilization of hydrocarbon feedstock in the preparation of low emission gasoline from cat cracker effluent. The method comprises (1) fractionating cat cracker effluent into a reduced emission gasoline blending stream, a heavy cat naphtha stream and a light cat recycle oil, (2) passing the heavy cat naphtha stream and optionally a portion of the light cat recycle oil to a selective membrane separation unit (3) recovering from the membrane separation unit an aromatics, sulfur and other heteroatoms permeate rich in saturated retentate, (4) passing the saturated rich retentate into the jet fuel or distillate depot, optionally following a hydrotreater treatment and (5) passing the aromatics / sulfur / hetero atoms rich permeate to a hydrocracker for conversion to hydrocracked material and gas.

Because of the restrictions placed on reduced emission gasoline in aromatics, polar materials and heteroatom content, heavy katnaphtha, which is traditionally blended in the mogas depot, should no longer be used. Valuable molecules would be lost or converted to a lower valuable product if the HCN were sent to the hydrocracker and converted into hydrocracked material and gas. The present process allows maximum recovery of valuable molecules from the HCN using a membrane separation, whereby a retentate suitable as a jet fuel or distillate is recovered after a choice of hydrotreating, while the detrimental aromatics, sulfur and other heteroatom molecules as a permeate suitable as a hydrocracker feed for conversion to hydrocracked material and gas.

Description of the Figures

Figure 1 is a typical current gasoline production schedule.

Figure 2 represents a method for producing reduced emission gasoline by excluding heavy catnaphtha from the mogas depot.

Figure 3 indicates the present invention wherein reduced emission gasoline produced by the valuable molecules in the HCN is recovered as a jet of high value fuel and distillate by membrane separation and hydrotreating.

The Present Invention

The production of gasoline with reduced emissions is currently the goal of environmentalists and governments with the aim of improving air quality.

The production of reduced emission gasoline requires that certain components currently present in gasoline such as aromatics, sulfur and other heteroatom molecules in the gasoline depot be limited.

Current gasoline production typically includes the use of certain fractions of cracker effluent as a gasoline feedstock.

The cracker effluent fractions are used because direct gasoline obtained by distillation of crude oil cannot meet demand. Cat cracker effluent contains significant amounts of hydrocarbons that boil in the gasoline range. Katkraker effluent is fractionated to recover the fraction boiling in that range. The cracker effluent is fractionated into fractions boiling in the naphtha range (about 232 ° C (450 ° F-)) and distillate boiling range (about 232 ° C + (450 ° F +)) also called light cat recycle oil. As currently used, fractionation of the cracker effluent produces a gasoline blending component containing aromatic hydrocarbons. These components are mainly the light and intermediate cat naphtha fractions mixed directly in the mogas depot and heavy cat naphtha and light cat recirculating oil. Part of the heavy katnaphtha fraction is mixed in the mogas depot because of its high aromatic content which contributes significantly to the octane level of the depot. However, the presence of aromatics is problematic in reduced emission gasoline.

Due to its content of aromatics, sulfur and other heteroatom molecules, this heavy catnaphtha will be excluded from the mogas depot in future mixing schedules.

Accordingly, the HCN together with the LCCO will typically be treated in a hydrocracker in the presence of hydrogen to form hydrocracked material and gas. The hydrocracked material can be mixed in the reduced emission mogas depot with or without intermediate processing, such as catalytic reforming.

However, this is not an economically attractive alternative because some of the more valuable hydrocarbon molecules present in the HCN are lost in the hydrocracking process and are converted into lower value hydrocracking material or as lower value gas. In addition, current hydrocracking capacity is not sufficient to absorb the additional volumes of HCN that become available for treatment once the use of HCN as a mogas blending component is prohibited. This would necessitate expensive new hydrocracking structures or an equally expensive hydrocracking expansion, which would also entail additional costs in ensuring the required amounts of hydrogen required by such expansions.

The present invention maximizes the utilization of hydrocarbon feedstock in the production of reduced emission gasoline by recovering the high value hydrocarbon components contained in the HCN and LCCO prior to hydrocracking.

In the present invention, the cracker effluent is still fractionated in light and intermediate cat naphtha suitable for direct mixing in the reduced gasoline mogas deposit and HCN emissions. The HCN, which boils in the 149 ° C + (300 ° F +) range (I49 * C-232eC) (300-450 ° F), which cannot be mixed directly in mogas along with a portion of the light cat recycle oil, is added to sent a selective membrane separation unit. This is accomplished not by physically mixing a portion of LCCO in the HCN, but by broadening the HCN cut-off point and thus to incorporate part of the LCCO into the HCN, the fractionator cut-off point is set higher so that the lower boiling part of it now classified as LCCO passes in the HCN, ie the naphtha fraction will boil in, for example, the 288eC (550 ° F); range, while the HCN, which is a wider cutoff fraction, will boil in the 149eC-288 ° C (300eF-550 * F) instead in the 149eC-232 "c (300-I50" F) range. The membrane unit divides the HCN / LCCO stream into an iromatics / sulfur / heteroatom-rich permeate and a saturated material-rich retentate.

The saturated material-rich retentate may optionally be at 150-100 ° C, preferably 250-350 ° C, 340-17200 KPa (50 to 2500 psig), preferably 580-3400 KPa (100-500 psig), 0 , 1 to 10 LHSV, preferably 1 to 5 LHSV, 8.4-252 m3H3 / m3 (50 to 1500 scf H3 / bbl), preferably 16.9-34.5 mH / m3 (100-500 scf H3 / bbl) hydrotreated above a catalyst active for hydrodesulfurization, hydrodenitrogenation, hydrodeoxygenation and / or aromatics and / or olefin hydrogenation, which catalysts are Ni / Mo, Co / Mo, Ni / Co / Mo, Co / W or Ni / W or may contain other transition metals on alumina, silica or zeolite supports. Pure or factory hydrogen can be used.

This saturated material-rich retentate (optionally hydrotreated) is suitable as a high value jet fuel blending material or as a distillate.

The permeate is sent to hydrocracking which is conducted at 200 to 500 ° C, preferably 300 to 400 ° C, 680-17200 KPa (100 to 2500 psi), preferably 3400-10200 KPa (500 to 1500 psig). 1 to 10 LHSV, preferably 1 to 5 LHSV and 84.5 to 1690 m3H2 / m3 (500 to 10,000 SCF Hj / bbl), preferably 338 to 1014 m'H / m3 (2000 to 6000 SCF H ^ bbl) above a catalyst selected from those cited above as suitable for the optional hydrotreating step, in either a single step or multi-step process. Pure or factory hydrogen can be used. The permeate is converted into hydrocracked material and gas.

An appropriate fraction of the hydrocracked material can be mixed in the REG mogas depot.

Optionally, the permeate for the preparation of a distillate product can be hydrogenated or used directly as a chemical feedstock.

The selective membrane separation zone operates under reverse osmosis, perstraction or pervaporation conditions, preferably pervaporation conditions.

The membrane separation zone allows the system described in II.S. U.S. Patent 3,370,102, which separates aromatics from saturated materials into a wide variety of feed materials including various petroleum fractions, naphthas, oils, and other hydrocarbon mixtures. Particularly mentioned in '102 is the separation of aromatics from kerosene. The process produces a permeate stream and a retentate stream, using a wiping liquid to remove the permeate from the front face of the membrane and thereby maintain the concentration gradient motive force. U.S. Patent 2,958,656 teaches the separation of hydrocarbons by type, ie aromatics, unsaturated, saturated by passing part of the mixture through a non-porous cellulose ether membrane and removing the permeate on the permeate side of the membrane by of a wipe gas or liquid. U.S. Patent 2,930,754 teaches a process for separating hydrocarbons of type, i.e., aromatics and / or olefins from gasoline boiling range blends, by the selective permeation of the aromatics through certain cellulose ester nonporous membranes. The permeated hydrocarbons are continuously removed from the permeate zone using a wipe gas or liquid. U.S. Patent 4,115,465 teaches the use of polyurethane membranes for the selective separation of aromatics from saturated materials via pervaporation.

U.S. Patent 4,914,064 describes polyurea / urethane membranes and their use for separating aromatics from non-aromatic hydrocarbons. The membrane is characterized in that it has a urea index of at least 20%, but less than 100%, an aromatic carbon content of at least 15 mol%, a functional group density of at least about 10 per 1000 grams of polymer and a C = 0 / NH ratio of less then has about 8.

Thin film combinations can be prepared either from slurry deposition as disclosed in U.S. Patent 4,861,628 or from solution deposition as disclosed in U.S. Patent 4,837,054.

The making of an anisotropic polyurea / urethane merabran is the subject of U.S. Patents 4,828,773 and 4,879,044. The anisotropic membrane has a three-layer structure, a thin dense layer formed on the film / support interface, a thin, non-continuous top layer formed at the membrane quench solvent interface, and an open, porous structure sandwiched between the aforementioned thin dense layer and the thin non-continuous top layer exists.

Polyurethane imides are produced by capping a polyol selected from those listed above with a polyisocyanate also selected from those listed above followed by chain extension by reaction with a polyanhydride directly producing the imide or with di- or polycarboxylic acids that produce amic acid groups that chemically or can be thermally condensed / cycled with the imide. Aliphatic and cycloaliphatic di- and polyisocyanates are useful, as well as mixtures of aliphatic, cycloaliphatic, aralkyl and aromatic polyisocyanates. Polyurethane imide membranes and their use for aromatics / non-aromatics separation are the subject of U.S. Patent 4,929,358.

Isocyanurate cross-linked polyurethane membranes and their use for the separation of aromatics from non-aromatics is the subject of U.S. Patent 4,929,357.

Acrylic ester homopolymers or their copolymers with each other or with acrylic acid are also useful for making aromatics / saturated material separation membranes. The acrylic acid monomer units can be in free acid form or partially or totally neutralized by metal or alkyl ammonium ions. The membranes can be covalently or ionically cross-linked.

Polyimide / aliphatic polyester copolymer membranes are examples of high temperature stable membranes that are preferably used in this separation.

Membranes of this type are described in U.S. Patents 4,944,880 and 4,990,275.

The polyester imide membranes are made from a copolymer comprising a polyimide segment and an oligomeric aliphatic polyester segment, the polyimide being derived from a dianhydride of between 8 and 20 carbon atoms and a diamine of between 2 and 30 carbon atoms and the oligomeric aliphatic polyester being a polyadipate, a polysucci -nate, a polymalonate, a polyoxalate or a polyglutarate.

The diamine is preferably diamine including phenylene diamine, methylenedianiline (MDA), methylene-di-o-chloroaniline (MOCA), methylene bis (dichloroaniline) (tetrachloro MDA), methylenedicyclohexyl amine (H12-MDA), methylenedichlorocyclohexylamine (HuM0CA), methylene-bis (dichlorocyclohexylamine) (tetrachloro H12MDA), 4,4 '- (hexafluoroisopropylidene) bisaniline (6F diamine), 3,3'-diaminophenylsulfone (3,3' DAPSON), 4,4 ' diaminophenyl sulfone (4,4'-DAPSON), 4,4'-dimethyl-3,3'-diaminophenyl sulfone (4,4'-dimethyl-3,3 'DAPSON), 2,4-diaminocumene, methyl bis (di- o-toluidine), oxydianiline (ODA), busaniline A, bisaniline M, bisaniline P, thiodianiline, 2,2-bis [4- (4-aminophenoxy) -phenyl] propane (BAPP), bis [4- (4 -aminophenoxyphenyl) sulfone (BAPS), 4,4'-bis (4-aminophenoxy) biphenyl (BAPB), 1,4-bis (4-aminophenoxy) benzene (TPE-Q) and 1,3-bis (4-aminophenoxy benzene (TPE-R).

The dianhydride is preferably an aromatic dianhydride and most preferably is selected from the group consisting of pyromellitic dianhydride, 3,3 ', 4,4'-benzophenone tetracarboxylic dianhydride, 4,4' - (hexafluoroisopropylidene) bis (phthalic anhydride), 4,4'-oxydiphthalic anhydride, diphenylsulfone-3,3 ', 4,4'-tetracarboxylic dianhydride and 3,3', 4,4'-bi-phenyl-tetracarboxylic acid anhydride.

Examples of preferred polyesters include polyethylene adipate, polyethylene succinate.

The polyesters used generally have molecular weights in the range of 500 to 4000, preferably 1000 to 2000.

In practice, the polyesterimide membrane can be synthesized as follows. One mole of a polyester polyadipate polysuccinate, polyoxalate, polyglutarate or polymalonate, preferably polyethylene adipate or polyethylene succinate, is reacted with two moles of the dianhydride, for example, pyromic acid dianhydride, to make a prepolymer in the end-capping step. One mole of this prepolymer is then added with one mole of diamine, e.g. methylene-di-o-chloroaniline (M0CA) converted to make a copolymer. Finally, heating the copolymer at 260-300 * 0 for about half an hour leads to the copolymer containing polyester and polyimide segments. Through the heating step, the polyamic acid is converted into the corresponding polyimide via the imide closure while removing water.

In the synthesis, an aprotic solvent such as dimethyl formamide (DMF) is used in the chain extension step. DMF is a preferred solvent, but other aprotic solvents are also suitable and useful. A concentrated solution of the polyamic acid / polyester copolymer in solvent is obtained. This solution is used for casting the membrane. The solution is spread on a glass plate or a high-temperature porous support, the layer thickness being adjusted by means of a casting knife. The membrane is first dried at room temperature to remove most of the solvent, then at 120 ° C overnight. If the membrane is cast on a glass plate, it is removed from the casting plate by soaking in water. When cast on a porous support it is left intact as such. Finally, heating the membrane at 300 ° C for about 0.5 hours results in the formation of the polyimide. Of course, heating to 300 ° C requires that if a support is used, the support is thermally stable, such as Teflon, fiber glass sintered metal or refractory material of a high temperature polymer backing.

As mentioned previously, the separation is accomplished under reverse osmosis, perstractive or pervaporation conditions.

In reverse osmosis, a membrane with pores in the range of 0.0001 to 1.0 / an, preferably 0.0001 to 0.001 fm is used and a sufficiently high applied pressure is applied to the osmotic pressure of one of the components in the feed to pass that component through the membrane. A typical reverse osmosis is performed at temperatures from O to 50 ° C at pressures up to 6800 KPa (1000 psig).

Perstraction involves the selective dissolution of certain components present in a mixture in the membrane, the diffusion of those components through the membrane and the removal of the diffused components on the downstream side of the membrane using a liquid wiping stream. In the abstract extraction of aromatics from saturates in petroleum or chemical streams (especially heavy katnaphtha streams), the aromatic molecules present in the feed stream dissolve in the membrane film by the similarity between the membrane solubility parameter and those of the aromatic compounds in the feed. The aromatics then pass through the membrane (diffuse) and are swept away by a wipe liquid with a low aromatics content. This keeps the concentration of aromatics on the permeate side of the membrane film low and maintains the concentration gradient responsible for permeation of the aromatics through the membrane.

The wiping liquid is preferably a saturated hydrocarbon liquid with a boiling point much lower or much higher than that of the aromatics passed. Provided the wiping liquid boiling range is much lower or higher than that of the aromatics passed, the presence of aromatics in that wiping liquid will not affect the permeate tomato concentration gradient (i.e. they will be in a different boiling range and therefore different aromatics). This will facilitate the separation, such as by simple distillation. Suitable wiping liquids therefore include, for example, C3-C6 saturated hydrocarbons and lubricating oil base materials (C15-Ca2).

The perstraction method is preferably carried out as low as possible at any suitable temperature.

The choice of pressure is not critical since the perstraction process does not depend on the pressure, but on the ability of the aromatic component in the feedstock to dissolve into and migrate through the membrane under a buoyancy concentration. Accordingly, any suitable pressure can be applied, the lower the better to avoid unwanted compaction if the membrane is supported on a porous backing, or tear the membrane if not so supported.

If C3 or Ct wiping liquids are used in the liquid state at 25 ° C or higher, the pressure must be increased to maintain them in the liquid phase.

Pervaporation is performed at generally higher temperatures than perstraction, where the feed material is either in liquid or vapor form and relies on vacuum or a wipe gas on the permeate side to evaporate or otherwise remove the permeate surface from the membrane and concentration gradient float -power that drives the separation process. As in perstraction, the aromatic molecules present in the feed dissolve in the membrane film, migrate through said film and reappear on the permeate side under the influence of a concentration gradient. Although pervaporation separation of aromatics from saturated materials can be performed at a temperature of about 25 ° C for the separation of benzene from hexane, the separation of heavier aromatics / saturated mixtures, such as heavy katnaphtha, requires higher temperatures of at least 80 ° C and above , preferably at least 100 ° C and above, most preferably 120 ° C and above (up to about 170 to 200 * 0 and above) are used, the maximum upper limit being that temperature at which the membrane will be physically damaged. vacuum of the order of 1-200 mm Hg is drawn on the permeate side The vacuum stream containing the permeate is cooled to allow the highly aromatic permeate to condense out The condensation temperature should be below the dew point of the permeate at a given vacuum level .

The membrane itself can take any suitable shape using any suitable module design. Thus, sheets of membrane material in a spiral wound or sheet and frame permeation cell module form can be used. Tubes and hollow fibers of membranes are useful in bundled configurations, where either the feed or the wiping fluid (or vacuum) is in the interior space of the tube or fiber, with the other material naturally on the opposite side.

Most suitably, the membrane is applied in traction in a hollow fiber configuration, where the feed is introduced on the outside of the fiber, the wiping liquid flows on the inside of the hollow fiber to wipe out the highly aromatic compound passed through, thereby the desired concentration gradient is maintained. The wiping liquid, along with the aromatics contained therein, is passed to a separator, typically a distillation means, however if a sufficiently low molecular weight wiping liquid is used, such as liquefied propane or butane, the wiping liquid can be easily evaporated recovering the liquid aromatics and recovering gaseous propane or butane (for example) and liquefying by applying pressure or lowering the temperature.

The practice of the present invention offers the following advantages.

In situations where the refiner does not have sufficient capacity to absorb the full HCN stream, the membrane separation can provide an alternative with a lower investment than a re-evaporation or new hydrocracker, especially if the existing hydrocrackers can be used as needed finishing the retate flow.

Total hydrogen consumption and hydrogen pressure conditions will be lower for the present process scheme. This is because the molecules branched from the hydrocracker by the membrane unit (ie the retentate) have been subjected to a lower pressure hydrotreating and also because the membrane separation has removed many of the sulfur and nitrogenous species from this stream (retentate) and those molecules in the permeate to be sent to the hydrocracker where they can be subjected to higher pressure hydrogenation.

The membrane hydrotreating process will result in overall higher yields of liquid hydrocarbons since the predominantly saturated retentate portion of the HCN-LCCO stream does not pass through the hydrocracker, in which some conversion to gas might take place.

The membrane hydrotreating system will result in a less overall deterioration in product value as earlier mogas HCN molecules are converted to hydrocracking material and some distillates instead of just hydrocracking material. Product improvement also occurs when low value LCCO components are included in the membrane feed for separation and yield higher value distillate products. In the future REG environment, where distillate margins are expected to rival those of mogas, these improvements and benefits will be even greater.

Figure 1 indicates the typical current process flow chart. Effluent from the cat cracker unit is sent via line 1 to a fractionation column (2) in which the effluent is split into light and intermediate catnaphtha (3) and heavy catnaphtha (4), which are mixed in the mogas depot (5). Light catabolic gas oil is sent via line (6) to a hydrocracker (7) for conversion to hydrocracking material (8) and gas (9).

Figure 2 presents a variation of the current typical process flow modified for the production of low emission gasoline. Cat cracker effluent is sent via line 1 to a fractionation column (2) where it is split into light and intermediate cat naphtha which is sent via line (3) to the reduced gasoline depot (4). Heavy catnaphtha boiling in a 149 ° C (300 ° F +) range (line 5), and light cat recycle oil (line 6) are sent to a hydrocracker (7) for conversion to gas (8) and hydrocracking material (9) .

Figure 3 illustrates the present invention.

Effluent from a cat cracker is sent via line 1 to a fractionation column 2 where it is split into light and intermediate cat naphtha sent via line 3 to the reduced emission gasoline depot 4. Heavy cat naphtha boiling at 149-288 ° C ( 300-450 ° F) range, or a wide fraction of a 149-288 ° C (300-550 ° F) range, is sent via line 5 to membrane unit 6. Light cat recycle oil is recovered from fractionation column 2 via line 7. LCCO is sent via line 7 to the hydrocracker 8. In membrane unit 6, the HCN is separated into a saturated material rich retentate which is sent via line 9 to the choice hydrotreater 10 for finishing to a jet fuel mix stock or distillate (11) or bypasses the hydrotreatment and is sent via line 9a directly to the jet fuel or distillate depot. A permeate (12) rich in aromatics, polar substances and heteroatom compounds from the membrane separation unit is sent to the hydrocracker 8 in which it is converted together with the LCCO from line 7 into hydrocracking material (13) and gas (14).

Claims (6)

A method for maximizing the utilization of a hydrocarbon feedstock in the production of reduced emission gasoline, the process comprising the following steps: (1) fractionating cat cracker effluent into a reduced emission gasoline blending stream, a heavy catnaphtha stream and a light katrec i reu1aato1i stream; (2) passing the heavy catnaphtha to a selective membrane separation unit; (3) recovering from the membrane separation unit a permeate rich in aromatics, sulfur and other heteroatoms and a retentate rich in saturated materials; (4) passing the saturated material rich retentate to the jet fuel or distillate depot; and (5) recovering the aromatics, sulfur and other heteroatom-rich permeate for further treatment. A method according to claim 1, characterized in that the saturated material-rich retentate recovered from the membrane separation unit is subjected to a hydrotreater treatment before being sent to the jet fuel or distillate depot. The method of claim 1, characterized in that the heavy catnaphtha fraction sent to the membrane separation unit is a wide fraction boiling in the range of 149-288 ° C (300 to 550 ° F). Process according to claim 1, characterized in that the aromatics, sulfur and other heteroatom-rich permeate for the preparation of hydrocracking material and gas is hydrocracked. Process according to claim 1, characterized in that the aromatics, sulfur and other heteroatom-rich permeate is hydrogenated to prepare a distillate or chemical feedstock. Process according to claim 1, characterized in that the membrane separation unit is operated under reverse osmosis, per-traction or pervaporation conditions.