PRODUCTION OF JET FUEL
United States Patent 3775291
Jet fuel, particularly suitable for use in supersonic aircraft, is produced from a mixture of a petroleum fraction boiling substantially in the kerosene range and a mixture of branched chain olefinic hydrocarbons having an average of nine to 16 carbon atoms per molecule. The mixture is passed through two hydrogenation zones in series, co-currently with hydrogen in the first, countercurrent to hydrogen in the second.
US Patent References:
/3594307.html
Kirk - July 1971 - 3594307
Mixed-phase hydrofining of hydrocarbon oils
Kelley et al. - September 1960 - 2952626
Two stage hydrogenation process
Hass et al. - September 1964 - 3147210
BLENDING HYDROGENATED FRACTIONS TO MAKE A JET FUEL
Barnes et al. - February 1970 - 3493491
PROCESS FOR MAKING JET FUEL
Barnes et al. - September 1970 - 3527693
/3594307.html
Kirk - July 1971 - 3594307
Mixed-phase hydrofining of hydrocarbon oils
Kelley et al. - September 1960 - 2952626
Two stage hydrogenation process
Hass et al. - September 1964 - 3147210
BLENDING HYDROGENATED FRACTIONS TO MAKE A JET FUEL
Barnes et al. - February 1970 - 3493491
PROCESS FOR MAKING JET FUEL
Barnes et al. - September 1970 - 3527693
Application Number:
05/177439
Publication Date:
11/27/1973
Filing Date:
09/02/1971
View Patent Images:
Export Citation:
Assignee:
The Lummus Company (Bloomfield, NJ)
Primary Class:
Other Classes:
208/143, 208/15
International Classes:
C10G65/08; C10G65/00; C10G23/00
Field of Search:
208/15,57
US Patent References:
Primary Examiner:
Levine, Herbert
Claims:
I claim
1. A process for producing jet fuel comprising the steps of:
2. A process according to claim 1 in which the mixture of
3. A process according to claim 2 in which the gas phase effluents from the first and second reaction zones are combined and passed in indirect heat exchange relationship with the feed to the first reaction zone, thereby cooling said gas-phase effluents and preheating said feed.
4. A process according to claim 1, wherein the gas-phase effluents from the first and second reaction zones are cooled sufficiently to condense the vaporized liquid components thereof, and said liquid components are separated from the remaining gas components and returned as liquid feed to the first reaction zone.
5. A process according to claim 4 wherein a major portion of said remaining gas components is returned as feed to the first reaction zone.
6. A process according to claim 4 wherein the volume ratio of recycled liquid to fresh feed is between 0.25:1.0 and 1.5:1.0.
7. A process according to claim 1 in which the petroleum fraction boiling substantially in the kerosene range is subjected to desulfurization prior to mixing in step (a).
8. A process according to claim 7 wherein the said petroleum fraction is a desulfurized straight-run kerosene.
9. A process according to claim 1 in which the petroleum fraction boiling substantially in the kerosene range. is mixed in step (a) with a mixture of branched chain olefinic hydrocarbons having an average of from 12 to 16 carbon atoms per molecule.
10. A process according to claim 9 in which the mixture of olefinic hydrocarbons comprises propylene tetramer.
11. A process according to claim 9 in which the mixture of olefinic hydrocarbons is an oligomer of butenes.
12. A process according to claim 1 wherein a liquid phase effluent is produced in step (e), which possesses a freezing point below -57°F.
13. A process according to claim 1 wherein the olefinic hydrocarbons are mixed in a ratio of 10 to 50 volume percent of the petroleum fraction.
14. A process according to claim 13 wherein the olefinic hydrocarbons are mixed in a ratio of 15 to 40 volume percent of the petroleum hydrocarbons.
15. A process according to claim 1 wherein the second hydrogenation zone is operated at a temperature of from about 300°F. to about 400°F.
1. A process for producing jet fuel comprising the steps of:
2. A process according to claim 1 in which the mixture of
3. A process according to claim 2 in which the gas phase effluents from the first and second reaction zones are combined and passed in indirect heat exchange relationship with the feed to the first reaction zone, thereby cooling said gas-phase effluents and preheating said feed.
4. A process according to claim 1, wherein the gas-phase effluents from the first and second reaction zones are cooled sufficiently to condense the vaporized liquid components thereof, and said liquid components are separated from the remaining gas components and returned as liquid feed to the first reaction zone.
5. A process according to claim 4 wherein a major portion of said remaining gas components is returned as feed to the first reaction zone.
6. A process according to claim 4 wherein the volume ratio of recycled liquid to fresh feed is between 0.25:1.0 and 1.5:1.0.
7. A process according to claim 1 in which the petroleum fraction boiling substantially in the kerosene range is subjected to desulfurization prior to mixing in step (a).
8. A process according to claim 7 wherein the said petroleum fraction is a desulfurized straight-run kerosene.
9. A process according to claim 1 in which the petroleum fraction boiling substantially in the kerosene range. is mixed in step (a) with a mixture of branched chain olefinic hydrocarbons having an average of from 12 to 16 carbon atoms per molecule.
10. A process according to claim 9 in which the mixture of olefinic hydrocarbons comprises propylene tetramer.
11. A process according to claim 9 in which the mixture of olefinic hydrocarbons is an oligomer of butenes.
12. A process according to claim 1 wherein a liquid phase effluent is produced in step (e), which possesses a freezing point below -57°F.
13. A process according to claim 1 wherein the olefinic hydrocarbons are mixed in a ratio of 10 to 50 volume percent of the petroleum fraction.
14. A process according to claim 13 wherein the olefinic hydrocarbons are mixed in a ratio of 15 to 40 volume percent of the petroleum hydrocarbons.
15. A process according to claim 1 wherein the second hydrogenation zone is operated at a temperature of from about 300°F. to about 400°F.
Description:
BACKGROUND OF THE INVENTION
This invention relates to the production of jet fuel from hydrocarbon feedstocks boiling substantially in the kerosene range. More particularly, this invention relates to a method for utilizing such hydrocarbons as feedstock for the production of jet fuel suitable for use in supersonic aircraft.
To be suitable for use in such aircraft, jet fuel must meet specifications exceeding in certain respects, those for jet fuel which can be used in ordinary, subsonic aircraft. Particularly, jet fuel, to be used for supersonic aircraft, must meet standards of low freezing point and high luminometer number. The freezing point of fuels, for use in such aircraft, must be at least as low as -57°F. or lower. See, for example, the specifications for "S.S.T. proposed fuel A" in Hydrocarbon Processing, Apr. 1971, Page 154, Table 2. The luminometer number is a measure of the burning characteristics of a fuel. The higher the number, the less smoke produced by the fuel during take-off and the lower the amount of radiation produced by the flame. This number is determined by the method described in ASTM designation D1322-54T. A high luminometer number is associated with low aromatics and also means at high smoke point.
It has been found that fuels having low luminometer numbers burn with a highly radiant flame and also produce excessive smoke during takeoff. Fuels having relatively high aromatic contents generally have relatively low luminometer numbers. It has also been found that the luminometer number can be increased generally by increasing the paraffin content of a jet fuel. It has further been found, as mentioned in U.S. Pat. No. 3,420,769 that the freezing point of a jet fuel can be lowered by isomerization of the straight chain paraffins to branched chained or isoparaffins.
Additionally, it has been found that materials boiling in the kerosene range provide suitable feedstocks for jet fuels. However, these materials, by themselves, even with treating in various manners, cannot be made to possess the low freezing points and the high luminometer numbers required of fuels for use in supersonic aircraft. Specifically, current specifications generally call for a fuel having a luminometer number in excess of 77 min. Additionally, such fuels will be required to meet other standards, such as high thermal stability, burning quality, and calorific value, particularly if they are to be considered suitable for supersonic transports.
It is an object therefore of this invention to produce jet fuels suitable for use in supersonic aircraft. It is a further object of this invention to produce jet fuels having a freezing point below -57°F. and a relatively high luminometer number. It is yet a further objective of this invention to provide such jet fuels with a low aromatics content. It is still a further object of this invention to provide such a jet fuel starting with a petroleum hydrocarbon boiling generally in the kerosene range. Other objects and advantages of this invention will be apparent from the description which follows.
It should be pointed out, however, that the fuels produced by this invention are not suitable for use only in supersonic aircraft. They are, in general, aviation turbine fuels of a high quality, which more than meet the specifications of fuels for use in jets flying below Mar. 1, as well as those which fly at supersonic speeds.
SUMMARY OF THE INVENTION
If brief, this invention contemplates the production of jet fuels by a process comprising:
a. mixing a petroleum fraction boiling substantially in the kerosene range with a mixture of branched-chain olefinic hydrocarbons having an average of from nine to 16 carbon atoms per molecule;
b. passing the resulting mixture in co-current contact with a hydrogen-rich gas through a first reaction zone operated at a temperature of from about 200°F. to about 575°F. and at an elevated pressure in contact with a hydrogenation catalyst;
c. removing from the first reaction zone a gas phase effluent comprising hydrogen and vaporized liquid materials, and a partially hydrogenated liquid effluent;
d. passing said liquid effluent into a second reaction zone operated at a temperature of from about 250°F. to about 450°F. and at an elevated pressure;
e. hydrogenating said liquid effluent by passing a hydrogen-rich gas into the second reaction zone countercurrently to it, in contact with a hydrogenation catalyst; and
f. drawing off from said second reaction zone a gas phase effluent comprising hydrogen and vaporized liquid materials and a liquid phase effluent comprising jet fuel.
DESCRIPTION OF THE DRAWING
The FIGURE is a diagrammatic illustration of the process of this invention.
DETAILED DESCRIPTION OF THE INVENTION
As shown in the FIGURE, the hydrogenation zones are preferably contained in one hydrogenation vessel, which has the form of a vertical cylinder having dished ends and pressure sustaining walls. The interior of the vessel is divided by horizontal partitions 12, 14, and 24, which are preferably perforated or foraminous plates or the like, into a plurality of chambers or zones including an upper reaction chamber 16, an intermediate vapor-disengaging zone 20, and a lower reaction chamber 18. The reaction chambers 16 and 18 are packed with a suitable hydrogenation catalyst 22, which may be of any of the well known hydrogenation-dehydrogenation catalysts, including such as Raney nickel, or nickel, platinum or palladium, preferably on a support such as alumina, silica, kieselguhr, diatomaceous earth, magnesia, zirconia or other inorganic oxides, alone or in combination. The catalyst in zone 16 is supported on partition 12. The catalyst in zone 18 is supported on a similar partition 24. Partition 24 is preferably spaced somewhat above the bottom of the converter, thus defining the upper boundary of an additional lower chamber or zone 26.
A fresh aromatics-containing feed boiling substantially in the kerosene range, as is hereinafter described, is introduced into the system at line 62. Into this stream, from line 60, is fed mixture of branched chain olefins, as hereinafter described, in an amount of 10-50 volume percent, preferably 15-40 percent, based on the aromatics-containing feed. The combined feed proceeds through line 46, and into it is introduced a hydrogen stream from line 38; it then proceeds through line 40 as indicated by the arrows, until it joins line 44, from which is added a condensed recycle liquid from separator 34. The resulting mixture of feed, recycle and hydrogen then passes through line 42 into the first zone 16 of the hydrogenation vessel which is operated at a temperature of from about 200°F. to about 500°F., and a pressure of from about 400 to about 1,500 psi.
The mixture of feed recycle liquid and hydrogen passes downwardly through the catalyst bed in zone 16, under adiabatic reaction conditions in which a substantial amount of the aromatics are hydrogenated to the corresponding naphthenic compounds, and substantially all of the olefins are also hydrogenated to the corresponding isoparaffins. The reaction product which passes out of zone 16 is a two-phase mixture. The liquid phase is a mixture of saturated and some unsaturated compounds. The gas phase effluent is a mixture of hydrogen, inert gaseous impurities, and vaporized liquid hydrocarbons of a composition generally similar to that of the liquid phase effluent.
The liquid phase of the effluent passes downwardly through the vapor-disengaging zone 20 into the second hydrogenation zone 18 (through partition 14, which serves as a distributor plate).
In reaction zone 18, hydrogen introduced through line 48 and passing through chamber 26 contacts the liquid phase effluent countercurrently, completing the hydrogenation of the aromatics. The hydrogen is introduced without being preheated, at a relatively low temperature, compared to that of the liquid phase effluent from zone 16, generally the hydrogen temperature is no higher than about 100°-120°F.
The liquid portion which emerges from hydrogenation zone 18 is briefly accumulated in chamber 26, permitting disengagement of the vapors and sealing the outlet to line 50 to prevent escape of hydrogen. The liquid portion is collected in line 50, and contains a very minor portion, generally less than 1 volume percent, of residual unhydrogenated aromatics, and virtually no olefins. The gas phase effluent from reaction zone 18 contains excess hydrogen, inert gaseous impurities and vaporized hydrocarbons similar to those produced in the gas phase effluent from zone 16.
The gas phase effluents from both the first hydrogenation zone 16 and the second hydrogenation zone 18 collect in vapor-disengaging zone 20. The combined gas phase fraction is withdrawn through line 28, and first passed through heat exchanger or waste heat boiler 52, in which some of the heat is used to produce steam for use in other processing steps, or in other processes, or for general purposes. The still hot vapor mixture is then passed through line 54, then preferably through condenser 32, where the vaporized liquid phase components remaining in the system are recondensed to liquids. The resulting two-phase system, consisting of gaseous hydrogen, inert gases, and reliquefied hydrocarbons, is passed into separator 34, where the liquid and gaseous phases are separated. The liquid phase is passed through line 44 to be mixed with the feed to hydrogenation zone 16 as previously described. The gaseous phase, comprising hydrogen and inert gases, may be vented partially, as through line 56, to prevent build-up of inert impurities in the system.
The remainder, and majority of this gaseous phase is recycled through line 36, to be mixed with the feed to the first hydrogenation zone 16 in line 40. Fresh feed hydrogen gas may be supplied from line 48 through line 58 into the recycle gas, in the event that the recycle hydrogen is insufficient to supply the needs in the first hydrogenation zone.
An important feature of this invention is a built-in temperature control. Reactions of the type contemplated are exothermic. The production of the desired jet fuel is favored by low outlet temperatures. Furthermore, runaway reactions must be prevented or coke and undesirable side products will be formed. Accordingly, external temperature control means are usually necessitated in processes for hydrogenating aromatics for jet fuel production. The present process, however, provides an inherent temperature control, particularly in the second hydrogenation zone 18. As the hydrogen feed from line 48 passes upwardly through this zone, a portion of the heat present in that chamber is absorbed in the process of sensibly heating the hydrogen. An additional amount of heat is absorbed by the vaporization of reaction product liquid in zone 18, in an amount sufficient to saturate the gas stream emerging from this zone into vapor-disengaging zone 20. Similarly, the temperature in the first reaction zone 16 is controlled by the absorption of heat in partially vaporizing the liquid feed. The vaporized liquid is removed from the vapor-disengaging zone 20 in conduit 28, as previously described. A similar process for the production of cyclohexane from benzene, with this same built-in temperature control, is described in U.S. Pat. No. 3,450,784.
The vaporized hydrocarbons recovered from the vapor disengaging zone 20 and used as recycle comprise partially hydrogenated feed containing up to about 5 percent aromatics. The volume ratio of recycle to fresh feed is generally in the range of about 0.25:1.0 to about 1.5:1.0, and depends on a number of factors, including hydrogen partial pressure and purity, desired temperature in the reactor, etc.
The major portion of the feed to this process comprises a petroleum fraction generally boiling in the kerosene range, particularly in the range from about 300°F. to about 550°F. This portion of the feed may be a straight run kerosene, heavy naphtha, furnace oil, catalytically cracked cycle oil, etc. The process of this invention does not accomplish desulfurization for practical purposes, though some desulfurization may take place; consequently most feeds should be desulfurized prior to being introduced into the process, generally in a separate unit (not shown).
The olefinic mixture utilized in the process of this invention is a medium molecular weight oligomer of propylene or of butene isomers, or a mixed oligomer of C 2 to C 4 olefins. The oligomer, being a polymer of olefins is itself unsaturated, and will contain nine to 16 carbon atoms per molecule. Most preferable are C 12 --C 16 olefinic compounds, with those having a greater degree of branching being especially preferred.
Generally, these oligomers and their mixtures boil in the range of about 280° to about 525°F., preferably 340° to about 500°F. Such oligomers as propylene trimer and tetramer, butene trimmers and tetramers, and mixed propylene-butene and ethylene-propylene-butene oligomers are suitable, as are mixtures such as polymer gasoline. Oligomers of butenes, or of butenes with ethylene and/or propylene, are expected to be particularly suitable, as these will display the greatest tendency toward forming branched chains.
If the kerosene portion of the feed has undergone desulfurization just prior to its admission into the first hydrogenation zone, it will be sufficiently hot that no further heating is required to bring it up to reaction temperature. If, however, it has been obtained from a simple fractionation process or has been allowed to cool down prior to being passed into this process, preheating is required. In any case, the hydrogen fed to the first hydrogenation zone 16 must be preheated prior to its introduction into this zone. The liquid recycle to this zone must also be preheated.
The hydrogenation of the olefinic oligomers is also an exothermic reaction, and occurs readily at somewhat lower temperatures than the hydrogenation of the kerosene - feed, which generally requires an inlet temperature of at least 350° to 380°F. Consequently, a saving in cost can be achieved, and the chance of overreaction minimized when the mixed feed is introduced into the reactor at a lower temperture, generally from about 200° to about 350°F. preferably 225°F. to 325°F. The maximum temperature leaving the first reaction zone generally is limited to 575°F. The exothermicity of the hydrogenation of the olefins will raise the temperature of the feed part way through the first reaction zone 16 and thus "trigger" the hydrogenation of the kerosene feed. Thus, if the kerosene feed does not require preheating, the olefinic feed can be mixed with it, without requiring preheating of the olefins.
The preheating of the hydrogen, recycle liquid, and feed if necessary, can be accomplished in a number of ways, and can be performed separately or in the same operation and equipment. A convenient method, in this process, is to utilize the heat contained in the vapors in lines 28 and 54, which have been removed from the vapor-disengaging zone 20. The combined hydrogen (and feed, if necessary) in stream 42, together with recycle liquid from line 44, is passed through heat exchanger 30, in which it is preheated to the desired inlet temperature by indirect heat exchange with the partially cooled vapors in line 54. This heat exchange, under some conditions, may have the additional effect of partially condensing some of the hydrocarbons in the combined vapor stream, facilitating the separation of hydrocarbons for recycle from the hydrogen and other gases, in separator 34.
If the fresh kerosene feed is already sufficiently hot so as not to require preheating, it should be by-passed around the preheater in a conventional manner to avoid overheating and undesirable side reactions. The fresh kerosene feed will then enter the system through line 43 instead of through line 46. Olefinic feed can be mixed from line 41. In this case, only the hydrogen and recycled liquid hydrocarbons will be preheated.
Alternatively, the preheating of the fresh feed, liquid recycle and hydrogen can be done in separate heat exchangers, and the heated materials mixed before being introduced into the reactor. This separate preheating can be done using any source of available heat, including the hot vapor mixture in line 54.
The ratio of hydrogen to fresh feed in the mixture fed to reaction zone 16 may vary from a minimum of the stoichiometric ratio of mol for each double bond to as much as about 3 mols for each double bond, and the ratio of hydrogen in line 48 to the liquid material entering reaction zone 18 may vary from about 0.3 to about 1.0 moles/mole.
The L.H.S.V. in the first zone 16 is preferably maintained between about 0.5 and about 6.0, based on fresh feed, while that in the second section 18 is generally at a higher level.
The temperature conditions in the second section should be adjusted to maintain the temperature of the liquid product at the outlet 50° between about 250°F and about 450°F, preferably between about 300°F. and about 400°F., to provide optimum conditions favoring hydrogenation of the aromatics to naphthenes and close equilibrium approach.
In order to illustrate more fully the nature of this invention, and the manner of practicing the same, the following specific example is presented.
EXAMPLE
A mixture of 80 percent straight run kerosene (which had been previously desulfurized and denitrified) and 20 percent untreated propylene tetramer was processed through the system shown in the FIGURE. The properties of the feeds were:
Kerosene Initial boiling point 348°F. End Point 500°F. Aromatics, volume % 16.5 Freezong point -52.5°F. Luminometer number, min. 48 Smoke point, min. 21.5 Propylene Tetramer Initial boiling point 358°F. End point 408°F. Freezing point below -73°F.
the Smoke Point of the mixed kerosene-propylene tetramer feed was 22.3 min and the luminometer number 50 min.
The reactor was operated at an inlet temperature (zone 16) of 417°F., a total pressure of 900 psig. and an overall L.H.S.V. of 1.83. The overall ratio of hydrogen to fresh feed hydrocarbons was 700 S.C.F. per barrel of feed.
The product of the reaction contained 1.3 volume percent of unsaturated aromatics, had a freezing point of -58°F., a Smoke Point of 34.6 mm. and a luminometer number of 86.
The above description is not intended to be all-inclusive as variations of the invention, according to the principles herein expressed, will no doubt readily occur to those skilled in the art. Accordingly, the invention shall not be deemed limited to the specific matters disclosed herein, but is only limited by the appended claims.
This invention relates to the production of jet fuel from hydrocarbon feedstocks boiling substantially in the kerosene range. More particularly, this invention relates to a method for utilizing such hydrocarbons as feedstock for the production of jet fuel suitable for use in supersonic aircraft.
To be suitable for use in such aircraft, jet fuel must meet specifications exceeding in certain respects, those for jet fuel which can be used in ordinary, subsonic aircraft. Particularly, jet fuel, to be used for supersonic aircraft, must meet standards of low freezing point and high luminometer number. The freezing point of fuels, for use in such aircraft, must be at least as low as -57°F. or lower. See, for example, the specifications for "S.S.T. proposed fuel A" in Hydrocarbon Processing, Apr. 1971, Page 154, Table 2. The luminometer number is a measure of the burning characteristics of a fuel. The higher the number, the less smoke produced by the fuel during take-off and the lower the amount of radiation produced by the flame. This number is determined by the method described in ASTM designation D1322-54T. A high luminometer number is associated with low aromatics and also means at high smoke point.
It has been found that fuels having low luminometer numbers burn with a highly radiant flame and also produce excessive smoke during takeoff. Fuels having relatively high aromatic contents generally have relatively low luminometer numbers. It has also been found that the luminometer number can be increased generally by increasing the paraffin content of a jet fuel. It has further been found, as mentioned in U.S. Pat. No. 3,420,769 that the freezing point of a jet fuel can be lowered by isomerization of the straight chain paraffins to branched chained or isoparaffins.
Additionally, it has been found that materials boiling in the kerosene range provide suitable feedstocks for jet fuels. However, these materials, by themselves, even with treating in various manners, cannot be made to possess the low freezing points and the high luminometer numbers required of fuels for use in supersonic aircraft. Specifically, current specifications generally call for a fuel having a luminometer number in excess of 77 min. Additionally, such fuels will be required to meet other standards, such as high thermal stability, burning quality, and calorific value, particularly if they are to be considered suitable for supersonic transports.
It is an object therefore of this invention to produce jet fuels suitable for use in supersonic aircraft. It is a further object of this invention to produce jet fuels having a freezing point below -57°F. and a relatively high luminometer number. It is yet a further objective of this invention to provide such jet fuels with a low aromatics content. It is still a further object of this invention to provide such a jet fuel starting with a petroleum hydrocarbon boiling generally in the kerosene range. Other objects and advantages of this invention will be apparent from the description which follows.
It should be pointed out, however, that the fuels produced by this invention are not suitable for use only in supersonic aircraft. They are, in general, aviation turbine fuels of a high quality, which more than meet the specifications of fuels for use in jets flying below Mar. 1, as well as those which fly at supersonic speeds.
SUMMARY OF THE INVENTION
If brief, this invention contemplates the production of jet fuels by a process comprising:
a. mixing a petroleum fraction boiling substantially in the kerosene range with a mixture of branched-chain olefinic hydrocarbons having an average of from nine to 16 carbon atoms per molecule;
b. passing the resulting mixture in co-current contact with a hydrogen-rich gas through a first reaction zone operated at a temperature of from about 200°F. to about 575°F. and at an elevated pressure in contact with a hydrogenation catalyst;
c. removing from the first reaction zone a gas phase effluent comprising hydrogen and vaporized liquid materials, and a partially hydrogenated liquid effluent;
d. passing said liquid effluent into a second reaction zone operated at a temperature of from about 250°F. to about 450°F. and at an elevated pressure;
e. hydrogenating said liquid effluent by passing a hydrogen-rich gas into the second reaction zone countercurrently to it, in contact with a hydrogenation catalyst; and
f. drawing off from said second reaction zone a gas phase effluent comprising hydrogen and vaporized liquid materials and a liquid phase effluent comprising jet fuel.
DESCRIPTION OF THE DRAWING
The FIGURE is a diagrammatic illustration of the process of this invention.
DETAILED DESCRIPTION OF THE INVENTION
As shown in the FIGURE, the hydrogenation zones are preferably contained in one hydrogenation vessel, which has the form of a vertical cylinder having dished ends and pressure sustaining walls. The interior of the vessel is divided by horizontal partitions 12, 14, and 24, which are preferably perforated or foraminous plates or the like, into a plurality of chambers or zones including an upper reaction chamber 16, an intermediate vapor-disengaging zone 20, and a lower reaction chamber 18. The reaction chambers 16 and 18 are packed with a suitable hydrogenation catalyst 22, which may be of any of the well known hydrogenation-dehydrogenation catalysts, including such as Raney nickel, or nickel, platinum or palladium, preferably on a support such as alumina, silica, kieselguhr, diatomaceous earth, magnesia, zirconia or other inorganic oxides, alone or in combination. The catalyst in zone 16 is supported on partition 12. The catalyst in zone 18 is supported on a similar partition 24. Partition 24 is preferably spaced somewhat above the bottom of the converter, thus defining the upper boundary of an additional lower chamber or zone 26.
A fresh aromatics-containing feed boiling substantially in the kerosene range, as is hereinafter described, is introduced into the system at line 62. Into this stream, from line 60, is fed mixture of branched chain olefins, as hereinafter described, in an amount of 10-50 volume percent, preferably 15-40 percent, based on the aromatics-containing feed. The combined feed proceeds through line 46, and into it is introduced a hydrogen stream from line 38; it then proceeds through line 40 as indicated by the arrows, until it joins line 44, from which is added a condensed recycle liquid from separator 34. The resulting mixture of feed, recycle and hydrogen then passes through line 42 into the first zone 16 of the hydrogenation vessel which is operated at a temperature of from about 200°F. to about 500°F., and a pressure of from about 400 to about 1,500 psi.
The mixture of feed recycle liquid and hydrogen passes downwardly through the catalyst bed in zone 16, under adiabatic reaction conditions in which a substantial amount of the aromatics are hydrogenated to the corresponding naphthenic compounds, and substantially all of the olefins are also hydrogenated to the corresponding isoparaffins. The reaction product which passes out of zone 16 is a two-phase mixture. The liquid phase is a mixture of saturated and some unsaturated compounds. The gas phase effluent is a mixture of hydrogen, inert gaseous impurities, and vaporized liquid hydrocarbons of a composition generally similar to that of the liquid phase effluent.
The liquid phase of the effluent passes downwardly through the vapor-disengaging zone 20 into the second hydrogenation zone 18 (through partition 14, which serves as a distributor plate).
In reaction zone 18, hydrogen introduced through line 48 and passing through chamber 26 contacts the liquid phase effluent countercurrently, completing the hydrogenation of the aromatics. The hydrogen is introduced without being preheated, at a relatively low temperature, compared to that of the liquid phase effluent from zone 16, generally the hydrogen temperature is no higher than about 100°-120°F.
The liquid portion which emerges from hydrogenation zone 18 is briefly accumulated in chamber 26, permitting disengagement of the vapors and sealing the outlet to line 50 to prevent escape of hydrogen. The liquid portion is collected in line 50, and contains a very minor portion, generally less than 1 volume percent, of residual unhydrogenated aromatics, and virtually no olefins. The gas phase effluent from reaction zone 18 contains excess hydrogen, inert gaseous impurities and vaporized hydrocarbons similar to those produced in the gas phase effluent from zone 16.
The gas phase effluents from both the first hydrogenation zone 16 and the second hydrogenation zone 18 collect in vapor-disengaging zone 20. The combined gas phase fraction is withdrawn through line 28, and first passed through heat exchanger or waste heat boiler 52, in which some of the heat is used to produce steam for use in other processing steps, or in other processes, or for general purposes. The still hot vapor mixture is then passed through line 54, then preferably through condenser 32, where the vaporized liquid phase components remaining in the system are recondensed to liquids. The resulting two-phase system, consisting of gaseous hydrogen, inert gases, and reliquefied hydrocarbons, is passed into separator 34, where the liquid and gaseous phases are separated. The liquid phase is passed through line 44 to be mixed with the feed to hydrogenation zone 16 as previously described. The gaseous phase, comprising hydrogen and inert gases, may be vented partially, as through line 56, to prevent build-up of inert impurities in the system.
The remainder, and majority of this gaseous phase is recycled through line 36, to be mixed with the feed to the first hydrogenation zone 16 in line 40. Fresh feed hydrogen gas may be supplied from line 48 through line 58 into the recycle gas, in the event that the recycle hydrogen is insufficient to supply the needs in the first hydrogenation zone.
An important feature of this invention is a built-in temperature control. Reactions of the type contemplated are exothermic. The production of the desired jet fuel is favored by low outlet temperatures. Furthermore, runaway reactions must be prevented or coke and undesirable side products will be formed. Accordingly, external temperature control means are usually necessitated in processes for hydrogenating aromatics for jet fuel production. The present process, however, provides an inherent temperature control, particularly in the second hydrogenation zone 18. As the hydrogen feed from line 48 passes upwardly through this zone, a portion of the heat present in that chamber is absorbed in the process of sensibly heating the hydrogen. An additional amount of heat is absorbed by the vaporization of reaction product liquid in zone 18, in an amount sufficient to saturate the gas stream emerging from this zone into vapor-disengaging zone 20. Similarly, the temperature in the first reaction zone 16 is controlled by the absorption of heat in partially vaporizing the liquid feed. The vaporized liquid is removed from the vapor-disengaging zone 20 in conduit 28, as previously described. A similar process for the production of cyclohexane from benzene, with this same built-in temperature control, is described in U.S. Pat. No. 3,450,784.
The vaporized hydrocarbons recovered from the vapor disengaging zone 20 and used as recycle comprise partially hydrogenated feed containing up to about 5 percent aromatics. The volume ratio of recycle to fresh feed is generally in the range of about 0.25:1.0 to about 1.5:1.0, and depends on a number of factors, including hydrogen partial pressure and purity, desired temperature in the reactor, etc.
The major portion of the feed to this process comprises a petroleum fraction generally boiling in the kerosene range, particularly in the range from about 300°F. to about 550°F. This portion of the feed may be a straight run kerosene, heavy naphtha, furnace oil, catalytically cracked cycle oil, etc. The process of this invention does not accomplish desulfurization for practical purposes, though some desulfurization may take place; consequently most feeds should be desulfurized prior to being introduced into the process, generally in a separate unit (not shown).
The olefinic mixture utilized in the process of this invention is a medium molecular weight oligomer of propylene or of butene isomers, or a mixed oligomer of C 2 to C 4 olefins. The oligomer, being a polymer of olefins is itself unsaturated, and will contain nine to 16 carbon atoms per molecule. Most preferable are C 12 --C 16 olefinic compounds, with those having a greater degree of branching being especially preferred.
Generally, these oligomers and their mixtures boil in the range of about 280° to about 525°F., preferably 340° to about 500°F. Such oligomers as propylene trimer and tetramer, butene trimmers and tetramers, and mixed propylene-butene and ethylene-propylene-butene oligomers are suitable, as are mixtures such as polymer gasoline. Oligomers of butenes, or of butenes with ethylene and/or propylene, are expected to be particularly suitable, as these will display the greatest tendency toward forming branched chains.
If the kerosene portion of the feed has undergone desulfurization just prior to its admission into the first hydrogenation zone, it will be sufficiently hot that no further heating is required to bring it up to reaction temperature. If, however, it has been obtained from a simple fractionation process or has been allowed to cool down prior to being passed into this process, preheating is required. In any case, the hydrogen fed to the first hydrogenation zone 16 must be preheated prior to its introduction into this zone. The liquid recycle to this zone must also be preheated.
The hydrogenation of the olefinic oligomers is also an exothermic reaction, and occurs readily at somewhat lower temperatures than the hydrogenation of the kerosene - feed, which generally requires an inlet temperature of at least 350° to 380°F. Consequently, a saving in cost can be achieved, and the chance of overreaction minimized when the mixed feed is introduced into the reactor at a lower temperture, generally from about 200° to about 350°F. preferably 225°F. to 325°F. The maximum temperature leaving the first reaction zone generally is limited to 575°F. The exothermicity of the hydrogenation of the olefins will raise the temperature of the feed part way through the first reaction zone 16 and thus "trigger" the hydrogenation of the kerosene feed. Thus, if the kerosene feed does not require preheating, the olefinic feed can be mixed with it, without requiring preheating of the olefins.
The preheating of the hydrogen, recycle liquid, and feed if necessary, can be accomplished in a number of ways, and can be performed separately or in the same operation and equipment. A convenient method, in this process, is to utilize the heat contained in the vapors in lines 28 and 54, which have been removed from the vapor-disengaging zone 20. The combined hydrogen (and feed, if necessary) in stream 42, together with recycle liquid from line 44, is passed through heat exchanger 30, in which it is preheated to the desired inlet temperature by indirect heat exchange with the partially cooled vapors in line 54. This heat exchange, under some conditions, may have the additional effect of partially condensing some of the hydrocarbons in the combined vapor stream, facilitating the separation of hydrocarbons for recycle from the hydrogen and other gases, in separator 34.
If the fresh kerosene feed is already sufficiently hot so as not to require preheating, it should be by-passed around the preheater in a conventional manner to avoid overheating and undesirable side reactions. The fresh kerosene feed will then enter the system through line 43 instead of through line 46. Olefinic feed can be mixed from line 41. In this case, only the hydrogen and recycled liquid hydrocarbons will be preheated.
Alternatively, the preheating of the fresh feed, liquid recycle and hydrogen can be done in separate heat exchangers, and the heated materials mixed before being introduced into the reactor. This separate preheating can be done using any source of available heat, including the hot vapor mixture in line 54.
The ratio of hydrogen to fresh feed in the mixture fed to reaction zone 16 may vary from a minimum of the stoichiometric ratio of mol for each double bond to as much as about 3 mols for each double bond, and the ratio of hydrogen in line 48 to the liquid material entering reaction zone 18 may vary from about 0.3 to about 1.0 moles/mole.
The L.H.S.V. in the first zone 16 is preferably maintained between about 0.5 and about 6.0, based on fresh feed, while that in the second section 18 is generally at a higher level.
The temperature conditions in the second section should be adjusted to maintain the temperature of the liquid product at the outlet 50° between about 250°F and about 450°F, preferably between about 300°F. and about 400°F., to provide optimum conditions favoring hydrogenation of the aromatics to naphthenes and close equilibrium approach.
In order to illustrate more fully the nature of this invention, and the manner of practicing the same, the following specific example is presented.
EXAMPLE
A mixture of 80 percent straight run kerosene (which had been previously desulfurized and denitrified) and 20 percent untreated propylene tetramer was processed through the system shown in the FIGURE. The properties of the feeds were:
Kerosene Initial boiling point 348°F. End Point 500°F. Aromatics, volume % 16.5 Freezong point -52.5°F. Luminometer number, min. 48 Smoke point, min. 21.5 Propylene Tetramer Initial boiling point 358°F. End point 408°F. Freezing point below -73°F.
the Smoke Point of the mixed kerosene-propylene tetramer feed was 22.3 min and the luminometer number 50 min.
The reactor was operated at an inlet temperature (zone 16) of 417°F., a total pressure of 900 psig. and an overall L.H.S.V. of 1.83. The overall ratio of hydrogen to fresh feed hydrocarbons was 700 S.C.F. per barrel of feed.
The product of the reaction contained 1.3 volume percent of unsaturated aromatics, had a freezing point of -58°F., a Smoke Point of 34.6 mm. and a luminometer number of 86.
The above description is not intended to be all-inclusive as variations of the invention, according to the principles herein expressed, will no doubt readily occur to those skilled in the art. Accordingly, the invention shall not be deemed limited to the specific matters disclosed herein, but is only limited by the appended claims.