EP1992673A1 - Isomerisierung von benzolhaltigen Rohmaterialien - Google Patents

Isomerisierung von benzolhaltigen Rohmaterialien Download PDF

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Publication number
EP1992673A1
EP1992673A1 EP08251358A EP08251358A EP1992673A1 EP 1992673 A1 EP1992673 A1 EP 1992673A1 EP 08251358 A EP08251358 A EP 08251358A EP 08251358 A EP08251358 A EP 08251358A EP 1992673 A1 EP1992673 A1 EP 1992673A1
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EP
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Prior art keywords
isomerization
conduit
hydrogenation
zone
benzene
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EP08251358A
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English (en)
French (fr)
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David James C/o UOP LLC Shecterle
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Honeywell UOP LLC
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UOP LLC
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Priority claimed from US11/750,521 external-priority patent/US7534925B2/en
Priority claimed from US11/750,523 external-priority patent/US20080286172A1/en
Application filed by UOP LLC filed Critical UOP LLC
Publication of EP1992673A1 publication Critical patent/EP1992673A1/de
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    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G45/00—Refining of hydrocarbon oils using hydrogen or hydrogen-generating compounds
    • C10G45/44—Hydrogenation of the aromatic hydrocarbons
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1037—Hydrocarbon fractions
    • C10G2300/104—Light gasoline having a boiling range of about 20 - 100 °C
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1037—Hydrocarbon fractions
    • C10G2300/1044—Heavy gasoline or naphtha having a boiling range of about 100 - 180 °C
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/10—Feedstock materials
    • C10G2300/1096—Aromatics or polyaromatics
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/20—Characteristics of the feedstock or the products
    • C10G2300/201—Impurities
    • C10G2300/202—Heteroatoms content, i.e. S, N, O, P
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4018—Spatial velocity, e.g. LHSV, WHSV
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2300/00—Aspects relating to hydrocarbon processing covered by groups C10G1/00 - C10G99/00
    • C10G2300/40—Characteristics of the process deviating from typical ways of processing
    • C10G2300/4081—Recycling aspects
    • C—CHEMISTRY; METALLURGY
    • C10—PETROLEUM, GAS OR COKE INDUSTRIES; TECHNICAL GASES CONTAINING CARBON MONOXIDE; FUELS; LUBRICANTS; PEAT
    • C10G—CRACKING HYDROCARBON OILS; PRODUCTION OF LIQUID HYDROCARBON MIXTURES, e.g. BY DESTRUCTIVE HYDROGENATION, OLIGOMERISATION, POLYMERISATION; RECOVERY OF HYDROCARBON OILS FROM OIL-SHALE, OIL-SAND, OR GASES; REFINING MIXTURES MAINLY CONSISTING OF HYDROCARBONS; REFORMING OF NAPHTHA; MINERAL WAXES
    • C10G2400/00—Products obtained by processes covered by groups C10G9/00 - C10G69/14
    • C10G2400/02—Gasoline

Definitions

  • This invention relates generally to the isomerization of hydrocarbons. This invention relates more specifically to the processing of benzene-containing hydrocarbon feeds and the isomerization of light paraffins.
  • High octane gasoline is required for modem gasoline engines. Benzene has a high octane number value and has been previously blended into gasoline. However, as benzene is phased out of gasoline for environmental reasons, it has become increasingly necessary to rearrange the structure of the hydrocarbons used in gasoline blending in order to achieve high octane ratings. Catalytic reforming and catalytic isomerization are two widely used processes for this upgrading.
  • a gasoline blending pool is usually derived from naphtha feedstocks and includes C 4 and heavier hydrocarbons having boiling points of less than 205°C (395°F) at atmospheric pressure.
  • This range of hydrocarbon includes C 4 -C 9 paraffins, cycloparaffins and aromatics.
  • C 5 and C 6 normal paraffins which have relatively low octane numbers.
  • the C 4 - C 6 hydrocarbons have the greatest susceptibility of octane improvement by lead addition and were formerly upgraded in this manner.
  • Octane improvement can also be obtained by catalytically isomerizing the paraffinic hydrocarbons to rearrange the structure of the paraffinic hydrocarbons into branch-chained paraffins or reforming to convert the C 6 and heavier hydrocarbons to aromatic compounds.
  • Normal C 5 hydrocarbons are not readily converted into aromatics, therefore, the common practice has been to isomerize these lighter hydrocarbons into corresponding branch-chained isoparaffins.
  • the non-cyclic C 6 and heavier hydrocarbons can be upgraded into aromatics through dehydrocyclization, the conversion of C 6 's to aromatics creates higher density species and increases gas yields with both effects leading to a reduction in liquid volume yields.
  • the benzene contribution from the reformate portion of the gasoline pool can be decreased or eliminated by altering the operation of the reforming section.
  • the operation of the reforming section may be altered to reduce the reformate benzene concentration. Changing the cut point of the naphtha feed split between the reforming and isomerization zones from 82 to 93°C (180° to 200°F) will remove benzene, cyclohexane and methylcyclopentane from the reformer feed.
  • Benzene can alternately also be removed from the reformate product by splitting the reformate into a heavy fraction and a light fraction that contains the majority of the benzene. Practicing either method will put a large quantity of benzene into the feed to the isomerization zone.
  • the isomerization of paraffins is a reversible reaction which is limited by thermodynamic equilibrium.
  • the basic types of catalyst systems that are used in effecting the reaction are a hydrochloric acid promoted aluminum chloride system and a supported aluminum chloride catalyst.
  • Either catalyst is very reactive and can generate undesirable side reactions such as disproporationation and cracking. These side reactions not only decrease the product yield but can form olefinic fragments that combine with the catalyst and shorten its life.
  • One commonly practiced method of controlling these undesired reactions has been to carry out the reaction in the presence of hydrogen. With the hydrogen that is normally present and the high reactivity of the catalyst, any benzene entering the isomerization zone is quickly hydrogenated. The hydrogenation of benzene in the isomerization zone increases the concentration of napthenic hydrocarbons in the isomerization zone.
  • Driers that are only operated at low pressures are less costly than high pressure driers and the cost of the many valves associated with the driers for the purposes of regenerating the drier sieves is reduced significantly for low pressure driers. Finally, additional utility savings are realized by the elimination of the condensing equipment normally required downstream of the hydrogenation reaction zone.
  • This invention is a process for converting a feedstock comprising C 4 -C 7 paraffins and C 5 -C 7 cyclic hydrocarbons including benzene.
  • This invention uses a hydrogenation zone upstream of the isomerization reactors to saturate benzene and simultaneously heat the feed to the isomerization zone.
  • the use of a separate hydrogenation zone also lowers the overall temperature of the isomerization zone feed as the benzene is saturated--lower temperatures minimize undesirable hydrocracking reactions.
  • Also performing the highly exothermic benzene saturation reaction in a lead reactor that has a lower temperature reduces the coking that could occur in the isomerization zone as a result of the higher overall temperatures.
  • this invention is a process for the isomerization of a C 4 -C 6 paraffinic feedstock that contains at least 1 wt.-% benzene.
  • the process includes the steps of combining the feedstock with a hydrogen-rich gas stream to produce a combined feed.
  • the combined feed is passed to a hydrogenation zone and contacted therein with a hydrogenation catalyst to saturate benzene and heat the feedstream.
  • the saturated feedstream is recovered from the hydrogenation zone and has a benzene concentration of less than 1.5 wt.-%.
  • At least a portion of the saturated feedstream is passed from the hydrogenation zone to an isomerization zone without heating and contacted with an isomerization catalyst at isomerization conditions.
  • this invention is a process for the isomerization of C 5 -C 6 paraffinic feedstock that contain at least 1 wt.-% benzene.
  • the process dries the feedstock before combining the feedstock with a dried hydrogen-rich gas to produce a combined feed that is passed at a temperature of from 38 to 232°C (100 to 450°F) to an hydrogenation zone and contacted therein with a hydrogenation catalyst.
  • the temperature of the combined feed is 127 to 232°C (260 to 450°F) or 149 to 204°C (300 to 400°F).
  • the combined feed may be heat exchanged with isomerization and hydrogenation zone effluents.
  • the saturated feedstream has a benzene concentration of from 0.01 to 5 wt.-% or from 0.1 to 1.5 wt.-% and is heat exchanged with the combined feed and the feedstock, and possibly cooled, before being passed to an isomerization zone.
  • the saturated feedstream is contacted with an isomerization catalyst in the isomerization zone to isomerize C 5 -C 6 hydrocarbons.
  • An isomerate product essentially free of benzene is recovered from the isomerization zone. Downstream separations may be used to recycle low octane components of the isomerization zone effluent.
  • the FIGURE shows a schematic flow diagram of one embodiment of the process.
  • FIGURE does not show all pumps, condensers, reboilers, instruments and other well-known items of processing equipment in order to simplify the drawing.
  • the discussion points out several items of traditional processing equipment that may be eliminated and thus provide both a capital cost savings and an operational cost savings.
  • a feedstream comprising at least C 5 and C 6 paraffins along with at least 1 wt.-% benzene enter the process through line 10 and pass through a sulfur guard bed 12 that removes sulfur from the feedstream.
  • the sulfur-depleted feedstream in line 13 is passed through a low pressure drier 11 to remove water. It is important to note that line 13 is not passed through a reactor, nor is hydrogen added, before being dried in low pressure drier 11. This eliminates the need for product condensers and a receiver on line 13.
  • the elimination of a commonly used condenser provides an operational and equipment cost savings, and the elimination of the receiver additionally eliminates the need for a hydrogen vent.
  • Hydrogen is a valuable component in refineries today, and conservation of hydrogen results in positive value for the refiner. Finally, only a low pressure drier 11 is required. High pressure driers and their associated valves are far more costly than low pressure driers and their associated valves. Thus a cost savings is realized in requiring only a low pressure drier as opposed to a high pressure drier.
  • Make-up hydrogen enters the process through line 14 and passes through a drier 16 for removal of water and sulfur.
  • the dried feedstream in line 15 and the dried hydrogen from line 17 are combined in line 18 to form a combined feed.
  • the combined feed 18 is heat exchanged in an exchanger 24 against the contents of line 20 which carries the effluent from a second isomerization reactor 22.
  • the contents of line 18 are further heat exchanged in a heat exchanger 26 against the contents of line 28 which carries the effluent from a first isomerization reactor 30.
  • the contents of line 18 are still further heat exchanged in a heat exchanger 25 against the contents of line 34 which carries the effluent from a dehydrogenation reactor 32.
  • the hydrogenation reactor 32 receives the contents of line 18, the combined feed.
  • the hydrogenation reactor saturates benzene present in the combined feed and further heats the combined feed.
  • Line 34 carries a saturated feed from hydrogenation reactor 32 to the first isomerization reactor 30.
  • a chloride-containing compound is injected into the contents of line 34 by a line 80.
  • a first stage of isomerization takes place in reactor 30. Following the first stage of isomerization, the effluent in line 28 is exchanged in heat exchanger 26 against the combined feed in line 18 as discussed above. Line 28 then carries the partially cooled isomerization effluent from reactor 30 to reactor 22. After further isomerization in reactor 22, an isomerate product is taken by line 20, heat exchanged against the combined feed in line 18 using heat exchanger 24 and then is passed to a fractionation column 38. Fractionation column 38 removes light gases from the isomerate product which are taken overhead by line 42 and withdrawn from the process through the top of a receiver 44 via line 50. Recycle is conducted back to fractionation column 38 via line 46.
  • the stabilized isomerate product is withdrawn from the bottom of fractionation column 38 by line 40.
  • stabilized isomerate product in line 40 is conducted to deisohexanizer 58 to separate low octane alkanes, such as normal or single branched isoparaffins and cyclic compounds such as cyclohexane, for recycle to the isomerization zone via line 64.
  • Valuable isomerate product in lines 60 and 62 are combined into final product 66.
  • Suitable feedstocks for this invention will include C 4 plus hydrocarbons up to an end boiling point of 250°C (482°F).
  • the feedstocks that are used in this invention will typically include hydrocarbon fractions rich in C 4 -C 6 normal paraffins.
  • the term "rich” is defined to mean a stream having more than 50% of the mentioned component.
  • the feedstock will include significant amounts of benzene.
  • the concentration of benzene in the feedstock will at least equal 1.0 wt.-% and will normally be higher.
  • the concentration of benzene may be from 1 to 25 wt.-%, and is expected to usually be in the range of 3 to 15 wt.-% or 5 to 12 wt.-%.
  • the other feed components will usually comprise C 5 -C 6 cyclic and paraffinic hydrocarbons with normal pentane, normal hexane, and isohexane providing most of the paraffinic components.
  • the benzene in one of the feeds may be much higher than 25 wt.-%. The dilution effect of combining the streams results in the benzene being at a manageable level.
  • the isomerization zone and hydrogenation zone catalysts are often sulfur sensitive. Suitable guard beds or adsorptive separation processes may be used to reduce the sulfur concentration of the feedstock.
  • the FIGURE shows the treatment of the feedstock to remove sulfur upstream of the feedstock drier, hydrogen addition point, and the hydrogenation zone. It is important that the sulfur guard bed be located upstream of the drier since water may be liberated from fresh guard bed adsorbent.
  • the feed stream is heated by heat exchange with the effluent of the benzene saturation reactor using a heat exchanger before being passed to the sulfur guard bed. If needed, additional heat may be input into stream 10 before reaching sulfur guard bed 12.
  • the feed stream may be heated with any suitable process stream such as the stabilizer bottoms or with a utility stream such as steam or hot oil.
  • the streams directed to the isomerization zone are first passed through at least one drier.
  • the drier for this purpose reduces water content to 0.1 ppm or less, and suitable adsorption processes for this purpose are well known in the art.
  • the specific placement of the driers in relation to the guard beds and other streams allows for low pressure driers to be used to dry the feedstream.
  • Low pressure driers and their associated regeneration switching valves are much less costly than high pressure driers and are less costly to operate as well.
  • the feedstock passes through a drier and the hydrogen stream passes through another drier before the feedstock and the hydrogen stream are combined to form the combined feed. It is important to note that both the feedstock drier and the hydrogen drier are upstream of the hydrogenation zone.
  • this flow scheme eliminates the need for sulfur guard beds on the hydrogen stream sent to the saturation reactor. All of the hydrogen used for the hydrogenation and isomerization zones is sent through hydrogen driers where both sulfur and water contained in the hydrogen stream are removed.
  • a hydrogen stream is combined with the feedstock to provide hydrogen for the hydrogenation and isomerization zones.
  • the hydrogen stream also undergoes drying or other treatment, such as sulfur removal, necessary for the sustained operation of the isomerization zone or hydrogenation zone.
  • the hydrogenation of benzene in the hydrogenation zone results in a net consumption of hydrogen.
  • hydrogen is not consumed by the isomerization reaction, the isomerization of the light paraffins is usually carried out in the presence of hydrogen. Therefore, the amount of hydrogen added to the feedstock should be sufficient for both the requirements of the hydrogenation zone and the isomerization zone.
  • the amount of hydrogen admixed with the feedstock varies widely.
  • the amount of hydrogen can vary to produce anywhere from a 0.01 to a 10 hydrogen to hydrocarbon ratio in the isomerization zone effluent. Consumption of hydrogen in the hydrogenation zone increases the required amount of hydrogen admixed with the feedstock.
  • the input through the hydrogenation zone usually requires a relatively high hydrogen to hydrocarbon ratio to provide the hydrogen that is consumed in the saturation reaction. Therefore, hydrogen will usually be mixed with the feedstock in an amount sufficient to create a combined feed having a hydrogen to hydrocarbon ratio of from 0.1 to 2. Lower hydrogen to hydrocarbon ratios in the combined feed are preferred to simplify the system and equipment associated with the addition of hydrogen.
  • the hydrogen to hydrocarbon ratio must supply the stoichiometric requirements for the hydrogenation zone.
  • an excess of hydrogen be provided with the combined feed.
  • the isomerization zone will have a net consumption of hydrogen often referred to as the stoichiometric hydrogen requirement which is associated with a number of side reactions that occur. These side reactions include saturation of olefins and aromatics, cracking and disproportionation. Due to the presence of the hydrogenation zone, little saturation of olefins and aromatics will occur in the isomerization zone.
  • the effluent from the hydrogenation zone should contain enough hydrogen to satisfy the hydrogen requirements for the isomerization zone.
  • the effluent from the hydrogenation zone has a hydrogen to hydrocarbon mole ratio of from 0.05 to 2, in another embodiment the ratio is 0.1 to 1.5 and in yet another embodiment the ratio is 0.1 to 1.0.
  • the combined feed in line 18 comprising hydrogen and the feedstock enter the hydrogenation zone.
  • the hydrogenation zone is designed to saturate benzene at relatively mild conditions.
  • the hydrogenation zone comprises a bed of catalyst for promoting the hydrogenation of benzene.
  • catalyst compositions include platinum group, tin or cobalt and molydenum metals on suitable refractory inorganic oxide supports such as alumina.
  • the alumina is an anhydrous gamma-alumina with a high degree of purity.
  • platinum group metals refers to noble metals excluding silver and gold which are selected from the group consisting of platinum, palladium, germanium, ruthenium, rhodium, osmium, and iridium.
  • Such catalysts have been found to provide satisfactory benzene saturation at conditions including temperatures as low as 38°C (100° F), pressures from 1400 to 4800 kPa(g) (200 to 700 psig), an inlet hydrogen to hydrocarbon ratio in the range of 0.1 to 2, and a 1 to 40 liquid hourly space velocity (LHSV).
  • Other suitable pressures include from 2068 to 4137 kPa(g) (300 to 600 psig) and from 2413 to 3792 kPa(g) (350 to 550 psig) and other suitable liquid hourly space velocities include from 4 to 20 and 8 to 20 hr -1 .
  • the feed entering the hydrogenation zone will be heated to a temperature in the range of 38 to 232°C (100 to 450°F), 127 to 232°C (260 to 450°F) or 149 to 204°C (300 to 400°F) by heat exchange with the effluent from the hydrogenation and isomerization zones.
  • the exothermic saturation reaction increases the heat of the combined feed and saturates essentially all of the benzene contained therein.
  • the effluent from the hydrogenation zone provides a saturated feed for the isomerization zone that will typically contain from 0.01 wt.-% to 5 wt.-% or from 0.1 wt.-% to 1.5 wt.-% benzene or from 0.1 to 1.0 wt.-% benzene.
  • the saturated feed from the hydrogenation reactor is typically at a temperature in the range of 149 to 288°C (200 to 550°F); 177 to 274°C (350 to 525°F); or 204 to 274°C (400 to 525°F).
  • the isomerization zone operates at a lower temperature range, so the heat of the saturated feed may be recovered and used to provide heat to other colder streams either within the process or from outside the process.
  • the saturated feed may be heat exchanged with the combined feed and with the feedstock. If the saturated feed is still too high in temperature even after heat exchange, the saturated feed may be cooled using conventional techniques.
  • the isomerization zone uses a solid isomerization catalyst to promote the isomerization reaction.
  • the zeolitic type isomerization catalysts are well known and are described in detail in US 3,442,794 and US 3,836,597 .
  • Other catalysts include those such as described in US 6,927,188 .
  • the high chloride catalyst on an alumina base that contains platinum is also well known in the art and not described in detail here.
  • This type of catalyst also contains a chloride component.
  • the chloride component termed in the art "a combined chloride" is present in an amount from 2 to 10 wt.-% based upon the dry support material.
  • the feedstock may be treated by any method that will remove water and sulfur compounds. Sulfur may be removed from the feedstock by hydrotreating. Adsorption processes for the removal of sulfur and water from hydrocarbon streams are also well known to those skilled in the art.
  • Inlet temperatures to, and temperatures within the reaction zone will usually range from 38°C to 260°C (100°F to 500°F) or 104°C to 204°C (220°F to 400°F) or 104°C to 177°C (220°F to 350°F). Lower reaction temperatures are preferred for purposes of isomerization conversion since they favor isoalkanes over normal alkanes in equilibrium mixtures.
  • the isoalkane product recovery can be increased by opening some of the cyclohexane rings produced by the saturation of the benzene.
  • maximizing ring opening usually requires temperatures in excess of those that are most favorable from an equilibrium standpoint.
  • temperatures in the range of 60° to 160°C are desired from a normal-isoalkane equilibrium standpoint but, in order to achieve significant opening of C 5 and C 6 cyclic hydrocarbon ring, the preferred temperature range for this invention lies between 100° to 200°C.
  • higher reaction temperatures are required to maintain catalyst activity.
  • the most suitable operating temperatures for ring opening and isoalkane equilibrium coincide and are in the range from 145° to 225°C.
  • the reaction zone may be maintained over a wide range of pressures. Pressure conditions in the isomerization of C 4 -C 6 paraffins range from 1380 kPa(g) to 4830 kPa(g) (200 to 700 psig). Higher pressures favor ring opening, therefore, embodiments may use pressures for this process in the range of from 2410 kPa(g) to 4830 kPa(g) (350 to 700 psig) when ring opening is desired.
  • the feed rate to the reaction zone can also vary over a wide range. These conditions include liquid hourly space velocities ranging from 0.5 to 12 hr -1 , or between 0.5 and 3 hr -1 .
  • operation of the reaction zone may also require the presence of a small amount of an organic chloride promoter.
  • the organic chloride promoter serves to maintain a high level of active chloride on the catalyst as small amounts of chloride are continuously stripped off the catalyst by the hydrocarbon feed.
  • the concentration of promoter in the reaction zone is usually maintained at from 30 to 300 ppm.
  • Suitable promoter compounds include oxygen-free decomposable organic chlorides such as perchloroethylene, carbon tetrachloride, proplydichloride, butylchloride, and chloroform to name only a few of such compounds.
  • chloride promoter after the hydrogenation reactor, as shown in the Figure, may be carried out at such a location to expose the promoter to the highest available temperature and assure its complete decomposition.
  • the need to keep the reactants dry is reinforced by the presence of the organic chloride compound which may convert, in part, to hydrogen chloride. As long as the process streams are kept dry, there will be no adverse effect from the presence of small amounts of hydrogen chloride.
  • a preferred manner of operating the process is in a two-reactor, reaction zone system.
  • the catalyst used in the process can be distributed equally or in varying proportions between the two reactors.
  • the use of two reaction zones permits a variation in the operating conditions between the two reaction zones to enhance isoalkane production.
  • the two reaction zones can also be used to perform cyclic hydrocarbon conversion in one reaction zone and normal paraffin isomerization in the other.
  • the first reaction zone can operate at higher temperature and pressure conditions that favor ring opening but performs only a portion of the normal to isoparaffin conversion.
  • the two stage heating of the combined feed e.g., as provided by exchangers 26 and 24, facilitates the use of higher temperatures therein in a first isomerization reactor.
  • the final reactor stage may operate at temperature conditions that are more favorable for isoalkane equilibrium.
  • Another benefit of using two reactors is that it allows partial replacement of the catalyst system without taking the isomerization unit off stream. For short periods of time, during which the replacement of catalyst may be necessary, the entire flow of reactants may be processed through only one reaction vessel while catalyst is replaced in the other.
  • the effluent of the process will enter separation facilities for the recovery of an isoalkane product.
  • the separation facilities divide the reaction zone effluent into a product stream comprising C 5 and heavier hydrocarbons and a gas stream which is made up of C 3 lighter hydrocarbons and hydrogen.
  • C 4 hydrocarbons are present, the acceptability of these hydrocarbons in the product stream will depend on the blending characteristics of the desired product, in particular vapor pressure considerations. Consequently, C 4 hydrocarbons may be recovered with the heavier isomerization products or withdrawn as part of the overhead or in an independent product stream.
  • Suitable designs for rectification columns and separator vessels to separate the isomerization zone effluent are well known to those skilled in the art.
  • the separation facilities can consist of a product separator and a stabilizer.
  • the product separator operates as a simple flash separator that produces a vapor stream rich in hydrogen with the remainder of its volume principally comprising C 1 and C 2 hydrocarbons.
  • the vapor stream serves primarily as a source of recycle hydrogen which is usually returned directly to the hydrogenation process.
  • the separator may contain packing or other liquid vapor separation devices to limit the carryover of hydrocarbons. The presence of C 1 and C 2 hydrocarbons in the vapor stream do not interfere with the isomerization process, therefore, some additional mass flow for these components is accepted in exchange for a simplified column design.
  • the remainder of the isomerization effluent leaves the separator as a liquid which is passed on to a stabilizer, typically a trayed column containing approximately 30 trays.
  • the column will ordinarily contain condensing and reboiler loops for the withdrawal of a light gas stream comprising at least a majority of the remaining C 3 hydrocarbons from the feed stream and a light bottoms stream comprising C 5 and heavier hydrocarbons.
  • the C 4 's are withdrawn with the light gas stream.
  • the light gas stream will ordinarily serve as fuel gas.
  • the stabilizer overhead liquid which represents the remainder of the isomerization zone effluent passes back to the fractionation zone as recycle input.
  • a C 5 plus naphtha fresh feed having a composition shown in the Table enters through line 10 and is heat exchanged with the hydrogenation zone effluent before being passed through sulfur guard bed 12 to remove sulfur components.
  • the sulfur-free feed is conducted in line 13 to low pressure drier 11 to remove water. Feed in line 13 may be combined with recycle normal alkanes in line 64 from deisohexanizer 58 prior to being dried in low pressure drier 11. Furthermore, if some or all of the feed is light reformate, it is expected that the light reformate will already be sulfur-free and sulfur guard bed 12 may be bypassed or eliminated.
  • Optional line 70 shown as a dashed line, shows light reformate feed bypassing sulfur guard bed 12. Reducing the amount of material passing through the sulfur guard bed may result in a smaller guard bed being required thus reducing costs.
  • Hydrogen in line 14 is dried in drier 16 and combined with dried feed in line 15 to form a combined feed.
  • Combined feed 18 is passed through a series of heat exchangers such as exchangers 24, 25 and 26 to heat the feed to a temperature of 149°C to 204°C (300° to 400° F) which then enters the hydrogenation reactor at a pressure of 3450 kPa(g) (500 psig).
  • the combined feed is contacted with a catalyst comprising a platinum metal on a chlorided platinum alumina support at an LHSV of 20.
  • Contact of the combined feed with the hydrogenation catalyst produces a saturated feed that is withdrawn by line 34 and has no more than 0.5 wt.-% benzene.
  • the hydrogenation zone heats the saturated feed to a temperature of 177 to 274°C (350° to 525° F).
  • the saturated feed is heat exchanged with the combined feed in line 18 and with the fresh feed in line 10. If necessary, the saturated feed may also be cooled.
  • the saturated feed in line 34 is passed on to the isomerization zone at a pressure of 3240 kPa(g) (470 psig).
  • Perchloroethylene is added to the saturated feedstream at a rate of 150 wt. ppm which then enters the reactor train 30 and 22 of the isomerization zone.
  • the saturated feed stream contacts an alumina catalyst such as one having 0.25 wt.-% platinum and 5.5 wt.-% chloride.
  • the converted isomerization zone feed passed out of the reactor train in line 20 at a temperature of 93 to 204°C (200 to 400°F) and a pressure of 3100 kPa(g) (450 psig) and has an exemplary composition as shown in the Table.
  • cooled isomerization zone effluent in line 36 enters the stabilizer column 38 for the recovery of the product and removal of light gases.
  • Column 38 has, for example, 30 trays and the feed may enter above tray 15.
  • Column 38 splits the isomerization zone effluent into an overhead 42 which is cooled and condensed 44 to provide a recycle 46 and a fuel gas stream 50. Because of the chloride in the stream, the fuel gas stream 50 is passed through scrubber 52 to remove any chloride and provide a scrubbed fuel gas stream 56. Spent caustic is removed from scrubber 52 in stream 54.
  • An isomerization zone product 40 is withdrawn from the bottom of stabilizer column 38 and has the exemplary composition shown in the Table.
  • Isomerization zone product 40 is passed to deisohexanizer 58 to separate low octane normal and monomethyl alkanes into stream 64 which may be recycled to combine with the feed stock in line 13.
  • the pentanes, dimethylbutanes, and some monomethyl alkanes removed in DIH overhead 60 are combined with the C6 naphthenes and C7+ in DIH bottoms 62 to form the process product stream 66.
  • This example demonstrates the ability of the process to saturate benzene using a flow scheme that allows low pressure feedstock driers and requires no condensing of the feed that would also require a receiver with hydrogen venting and additional pumps.
  • the combined feed is heat exchanged with the effluents of the isomerization reactors and the benzene saturation reactor, and the benzene saturation effluent is also heat exchanged with the fresh feed. All values in the table are merely exemplary of one embodiment, and the compositions of the stream may vary with different applications.

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  • Chemical & Material Sciences (AREA)
  • Oil, Petroleum & Natural Gas (AREA)
  • Engineering & Computer Science (AREA)
  • Chemical Kinetics & Catalysis (AREA)
  • General Chemical & Material Sciences (AREA)
  • Organic Chemistry (AREA)
  • Production Of Liquid Hydrocarbon Mixture For Refining Petroleum (AREA)
  • Organic Low-Molecular-Weight Compounds And Preparation Thereof (AREA)
EP08251358A 2007-05-18 2008-04-08 Isomerisierung von benzolhaltigen Rohmaterialien Withdrawn EP1992673A1 (de)

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US11/750,521 US7534925B2 (en) 2007-05-18 2007-05-18 Isomerization of benzene-containing feedstocks
US11/750,523 US20080286172A1 (en) 2007-05-18 2007-05-18 Isomerization of Benzene-Containing Feedstocks

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WO2010151452A2 (en) 2009-06-25 2010-12-29 Uop Llc Light paraffin isomerization with improved feed purification
WO2014060462A2 (de) 2012-10-18 2014-04-24 Basf Se Kohlenwasserstoffkonversionsverfahren in gegenwart einer sauren ionischen flüssigkeit mit vorgeschalteter hydrierung
WO2014060461A1 (de) 2012-10-18 2014-04-24 Basf Se Neues verfahren zur herstellung von cyclohexan aus methylcyclopentan und benzol
WO2014060460A2 (de) 2012-10-18 2014-04-24 Basf Se Verfahren zur herstellung von cyclohexan mit aus einem steamcrackverfahren stammenden ausgangsmaterialien
EP2931685A4 (de) * 2012-12-14 2016-07-27 Uop Llc Verfahren und vorrichtungen für erhöhte alkyl-cyclopentan-konzentrationen in aromastoffreichen strömen
US9873646B2 (en) 2014-04-22 2018-01-23 Basf Se Process for preparing cyclohexane from benzene and methylcyclopentane with upstream benzene hydrogenation
US10081580B2 (en) 2012-10-18 2018-09-25 Basf Se Process for preparing cyclohexane with starting materials originating from a steamcracking process
US10351788B1 (en) 2018-02-28 2019-07-16 Uop Llc Processes and apparatus for isomerizing hydrocarbons
CN112552963A (zh) * 2019-09-10 2021-03-26 南京延长反应技术研究院有限公司 一种煤直接液化的智能控制强化系统及工艺
US11697777B2 (en) 2019-08-02 2023-07-11 Abu Dhabi Oil Refining Company—Takreer Single reactor process for benzene-saturation/isomertzation of light reformates

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US20140171706A1 (en) * 2012-12-14 2014-06-19 Uop Llc Methods and apparatuses for forming low-aromatic high-octane product streams
RU199611U1 (ru) * 2020-06-15 2020-09-09 Юрий Николаевич Киташов Реактор изомеризации дистиллята

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EP2445631A2 (de) * 2009-06-25 2012-05-02 Uop Llc Isomerisierung leichter paraffine mit verbesserter rohstoffreinigung
EP2445631A4 (de) * 2009-06-25 2014-09-10 Uop Llc Isomerisierung leichter paraffine mit verbesserter rohstoffreinigung
WO2010151452A2 (en) 2009-06-25 2010-12-29 Uop Llc Light paraffin isomerization with improved feed purification
US10081580B2 (en) 2012-10-18 2018-09-25 Basf Se Process for preparing cyclohexane with starting materials originating from a steamcracking process
WO2014060462A2 (de) 2012-10-18 2014-04-24 Basf Se Kohlenwasserstoffkonversionsverfahren in gegenwart einer sauren ionischen flüssigkeit mit vorgeschalteter hydrierung
WO2014060461A1 (de) 2012-10-18 2014-04-24 Basf Se Neues verfahren zur herstellung von cyclohexan aus methylcyclopentan und benzol
WO2014060460A2 (de) 2012-10-18 2014-04-24 Basf Se Verfahren zur herstellung von cyclohexan mit aus einem steamcrackverfahren stammenden ausgangsmaterialien
EP2931685A4 (de) * 2012-12-14 2016-07-27 Uop Llc Verfahren und vorrichtungen für erhöhte alkyl-cyclopentan-konzentrationen in aromastoffreichen strömen
US9873646B2 (en) 2014-04-22 2018-01-23 Basf Se Process for preparing cyclohexane from benzene and methylcyclopentane with upstream benzene hydrogenation
US10351788B1 (en) 2018-02-28 2019-07-16 Uop Llc Processes and apparatus for isomerizing hydrocarbons
WO2019169099A1 (en) * 2018-02-28 2019-09-06 Uop Llc Processes and apparatus for isomerizing hydrocarbons
CN111936453A (zh) * 2018-02-28 2020-11-13 环球油品有限责任公司 用于异构化烃的方法和设备
CN111936453B (zh) * 2018-02-28 2023-05-05 环球油品有限责任公司 用于异构化烃的方法和设备
US11697777B2 (en) 2019-08-02 2023-07-11 Abu Dhabi Oil Refining Company—Takreer Single reactor process for benzene-saturation/isomertzation of light reformates
CN112552963A (zh) * 2019-09-10 2021-03-26 南京延长反应技术研究院有限公司 一种煤直接液化的智能控制强化系统及工艺

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CA2628361C (en) 2012-08-07
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