US 3224959 A
Abstract available in
Claims available in
Description (OCR text may contain errors)
V DeC- 21, 1955 w G. scHLlNGER ETAL 3,224,959
HYDROCON'V'ERSION OF HYDROCARBONS WITH THE USE OF A TUBULAR REACTOR IN THE PRESENCE OF HYDROGEN AND THE RECYCLING OF A PORTION OF THE TAR-LIKE VISCOUS RESIDUE Filed Aug. '7. 1962 United States Patent O M 3,224,959 HYDRUCNVERHUN OF HYDROCAREONS WITH THE USE F A TUBULAR REACTR llN THE PRESENCE @lli HYDRGEN AND THE RE- CYCMNG F A PRTHN 0F ifi-lli TAR-LIKE JlSCUUS RESlDUE Warren G. Schlingen Pasadena, Charles H. Brodeur, Claremont, and Charles l?. Marion, Whittier, Calif., assiguors to rEeraco lne., New York, NSY., a corporation of Delaware Filed Aug. 7, i962, Ser. No. 215,379 lli Claims. (Cl. 20S-107) This invention relates to the treatment of hydrocarbons. More particularly, this invention is concerned with the conversion of heavy hydrocarbon liquids having high carbon and/or metal contents into lighter hydrocarbon oils of reduced carbon and/or metallic content in the presence of hydrogen.
Several methods for the hydroconversion of heavy hydrocarbons into lighter hydrocarbons are known. One particularly advantageous method is presented in copending US. patent application Serial No. 33,582, filed June 2, 1960, now U.S. Patent No. 3,089,843, of which one of us is co-inventor. In said application, there is disclosed a process for the conversion of hydrocarbon oils in which the hyd-rocarbon oil, in intimate mixture with hydrogen, is passed through a tubular reaction zone at elevated temperature and pressure under conditions of highly turbulent ilow. The effluent from the tubular `reaction zone is then introduced into a separation zone wherein the gaseous material is separated from material which is liquid at the prevailing conditions. The separated liquid material then, at substantially the same temperature and pressure, is contacted with a separately heated stream of hydrogen. Some additional hydroconversion takes place in the contacting zone together with removal of entrained or dissolved gaseous material, the resulting gaseous stream comprising unreacted hydrogen and vaporous hydrocarbons is combined with the gaseous material removed from the tubular reaction zone etliuent and the combined stream is then passed into contact with a hydrogenation catalyst.
The liquid bottoms removed from the contacting zone are tar-like in nature in that they have a high carbon residue, a high pour point and a high viscosity. The bottoms usually also have high sulfur and nitrogen contents. ln addition all of the metals present in the feed are concentrated in the bottoms. In many instances, depending on the feed and/or operating conditions this material is solid at room temperature.
Because of the tar-like nature of the contacting zone bottoms, it was not considered advisable to subject this material to further treatment to upgrade it into desired liquid products. Ordinarily, heavy hydrocarbon materials may be upgraded as by the-rmal cracking or colting. Such procedures are not too satisfactory however as the products for the most part are coke and xed gases and the yields of the desired liquid products are relatively small. ln the case of the bottoms from the contacting zone even such treatment was not considered practical because of the tar-like nature of this material. lt was believed from past experience that subjecting this tarry material to conventional treatment would produce large amounts of coke with negligible yields of middle distillates. Accordingly, it Was considered that the most expedient method of utilizing this bottoms product was to add thereto a light cutter oil to reduce its viscosity thereby producing a pumpable mixture which could be used as a fuel of the Bunker C type.
3,224,959 Patented Dec. 2l., 1965 ICC According to the present invention the bottoms from the contacting zone are upgraded to good. yields of desired liquid products by being recycled and introduced with fresh feed into the tubular reaction zone, Where, contrary to expectations, good yields of middle distillates are obtained.
The novel process is also characterized by less of the fresh feed going to tar despite the greatly increased carbon residue of the mixed feed. Another feature of the invention is that the heavy tar-like bottoms from the contacting zone are upgraded to good yields of middle distillates with the production of negligible amounts of coke. Another feature of the invention is that the greater conversion to desired liquid products is not accompanied by proportionately greater hydrogen consumption. It would appear that although all of the heavy components of the fresh feed are not converted to lighter material in the first pass through they tubular reactor, some change does take place and the bottoms are not a heavy, refractory diilculty-convertible tar-like residue but instead are as susceptible to conversion as the fresh feed. As a result, the bottoms can be recycled in large amounts to provide good yields of middle distillates Without causing the reactor to become plugged by the formation of coke.
Any hydrocarbon liquid may be used as the fresh feed to the process of the present invention. In this respect, the term fresh feed identies hydrocarbon material which is being introduced to the tubular reaction zone for the first time, as distinguished from lrecycle material. However, the process has particular application in the treatment of hydrocarbon liquids containing residual com ponents, metals and other tar and ash forming constituents. Hydrocarbon liquids which are particularly suitable as feed stocks for the process of the present invention are those having Conradson carbon values of at least about 1% by Weight. Examples of charge stocks to which the process of the invention may be applied successfully are crude oils such as Santa Maria crude, San Ardo crude, Arabian crude and heavy fractions of crude oils such as reduced or topped crude, deasphalted oil, vacuum residuurn and mixtures thereof and the like. Other materials which may be advantageously treated are coil oil, pitches, tars, gilsonite, shale oil and ta sand oil.
The hydrogen employed in the process of the present invention may be substantially pure, eg. SiO-99% by volume or may be dilute hydrogen such as a gas mixture containing as little as 40% by volume hydrogen obtained for example by the partial combustion of carbonaceous fuels. Suitable sources of hydrogen are catalytic reformer hydrogen, electrolytic hydrogen or synthesis gas which last may be used as produced or which may be used after being treated with a Water gas shift conversion catalyst and then scrubbed `for C()2 removal. The term hydrogen as used in the present specification and appended claims includes not only pure hydrogen but also includes dilute hydrogen. Preferably, the gas referred to as hydrogen contains at least about volume percent hydrogen.
The hydrogen and the hydrocarbon pass through the tubular reaction zone under such conditions of temperature, pressure and turbulence that the reactants are in the form of an intimate mixture. The turbulence level, that is the ratio of the average apparent viscosity to the kinematic viscosity, should be at least 25. In actual practice, the turbulence level is usually much greater, generally in excess of 100. Under these conditions the oil or at least that portion which is liquid under the prevailing conditions is in the form of ne inist-like droplets suspended in a gaseous medium comprising hydrogen, Under these conditions, the hydrogen is in close proximity to any cracked fragments which are formed during the hydroconversion so that the unsaturated cracked fragments can react with the hydrogen in preference to interreacting to form larger hydrocarbon molecules.
According to the first step of the process of the present invention, the fresh feed, the contacting or stripping zone bottoms and the hydrogen are introduced at elevated temperature and pressure into a tubular reaction zone through which the reaetant stream is passed under conditions of highly turbulent flow. Ordinarily the temperature in the tubular reaction Zone is maintained between about 700 and 1000 F., preferably 800 to 950 F. Pressure in the reaction zone is advantageously maintained between about 500 and 5000 p.s.i.g. Economically satisfactory results are obtained when the outlet pressure of the tubular reaction zone is between about 1000 and 2000 p.s.i.g. A large excess of hydrogen over the stoichiometric amount is circulated. Hydrogen rates of between about 1000 and 95,000 s.c.f.b. (standard cubic feet per barrel) of fresh feed are used, preferred rates being between 3,000 and 15,000 s.c.f.b.
The effluent from the tubular reaction zone in the form of finely-divided oil droplets suspended in a vaporous medium comprising hydrogen and vaporous hydrocarbons, is passed into a separation zone wherein the vaporous materials are separated from the liquid materials. The separated liquid material is then contacted at substantially the same temperature and pressure conditions as those prevailing in the tubular reaction zone with a separately heated hydrogen stream in an amount ranging between about 4,000 and 95,000 s.c.f.b. of fresh feed preferably 5,000 to 25,000 s.c.f.b.
During the contacting or stripping, some additional hydroconversion, as is evidenced by a slight increase in temperature, takes place in the contacting or stripping zone. In addition entrained or dissolved gaseous rnaterials are removed from the liquid and combined with the overhead from the separation Zone. The treated liquid material which has a relatively high carbon and metal content with respect to the fresh feed or at least a portion thereof is then recycled to the tubular reaction zone.
For a better understanding of the invention, reference is made to the accompanying drawing which illustrates diagrammatically a preferred embodiment of the present invention.
Charge oil (as described above) in line 11 is mixed with a large excess of hydrogen from line 21 under elevated pressure and the mixture is introduced into heater 13 wherein it is passed through a coil heated indirectly as by oil or gas combustion. The hydrogen is mixed with the oil in an amount ranging from 1000 to 95,000 s.c.f.b. of liquid feed, preferably from about 3,000 to about 15,000 s.c.f.b. of oil. Reaction temperatures within the coil are maintained between about 700 and 1000o F., preferably between about 800 and 950 F. The heating coil outlet pressure is advantageously maintained within the range of 1000 to 2000 p.s.i.g., although pressures ranging from as low as 500 p.s.i.g. to as high as 5000 p.s.i.g. or higher may be employed.
The ow rate is such as to keep the reactant mixture in a state of extreme turbulence. In this state, the higher boiling hydrocarbons, under the conditions of temperature, hydrogen to oil ratio, contact time and pressure, are subjected to viscosity breaking with substantially immediate hydrogenation of the molecular fragments and without further breakdown, thereby materially increasing the production of middle distillates Aboiling in the 400- 700" F. range without substantial increase in lower boiling gasoline range materials and without substantial formation of normally gaseous hydrocarbons and coke. As
the reaction proceeds, the molecular fragments, because of their lower boiling characteristics, are substantially immediately vaporized.
The hot mixture of hydrogen vaporized hydrocarbons and suspended liquid oil leaves heater 13 by means of line 15 and is introduced substantially immediately into the upper section of tower 17. This section serves as a disengaging zone to separate gasiform materials comprising vaporous hydrocarbons and hydrogen from the liquid oil. The gasiform materials leave tower 17 through line 3l. and the disengaged liquid flows downwardly through tower 17.
A separate hydrogen stream amounting to between 4000 and 95,000 s.c.f.b. of fresh feed to heater 13, and preferably between about 5,000 and 25,000 s.c.f.b. of fresh feed, is passed through line 19 to heater 27 wherein the hydrogen is heated indirectly as by oil or gas combustion to an elevated temperature between 800 and l000 F., preferably between about 860 and 925 F., while maintaining the outlet pressure of heater 27 substantially the same as the outlet pressure of heater 13. The hot hydrogen passes continuously from heater 27 through line 29 to the lower section of tower 17 at a point spaced above the bottom of the tower. The pressure within tower 17 is maintained at approximately the pressure of lines 15 and 29, preferably within the range of 1000 to 2000 pounds per square inch, The disengaged liquid oil which has been separated from the vaporous hydrocarbons and hydrogen in the upper section of tower 17 Hows downwardly and during its passage to the bottom of the tower is brought into intimate contact with the separately heated stream of hydrogen entering the lower section of tower 17 through line 29. Although the main action in the tower is a stripping of the lighter components from the liquid by the large quantity of contacting hydrogen, under the existing conditions of temperature, pressure and ratio of hydrogen to oil some further cracking and viscosity breaking takes place with hydrogenation of the unsaturated fragments. The combined effect is the production of additional oil vapors and the reduction of the liquid oil to a heavy residuum containing substantially all of the metal and ash-forming constituents of the original charge oil. The vaporized portion of the downwardly flowing oil is swept upwardly by the reaction hydrogen and passes through the tower 17 and out of the upper section thereof through line 31 with the vaporous hydrocarbons and hydrogen separated from liquid oil in the disengaging zone.
The remaining liquid hydrocarbons which contain substantially all of the metal contaminants present in the original hydrocarbon charge stock ow downwardly through tower 17 below the point of entry of line 29 and collect as a residual liquid fraction in boot 34 from which they may be withdrawn continuously or intermittently.
The yield of bottoms collected in boot 34 may be controlled as desired depending on the character of the oil charge to the heater-reactor 13, by varying the charge rate, the temperature and/ or pressure in the heating coil, the temperature within the tower and the relative quantity of hydrogen introduced through lines 1S and 29. The highest yield of treated vaporized oil which passes overhead through line 31 and the lowest yield of unvaporized oil collected in boot 34 is obtained by operating at increased temperature 'and time of reaction and increased quantity of hydrogen charged. In general, it is most desirable from an economic standpoint to carry out the operation under conditions to produce the lowest yield of bottoms while conducting the operation `at such a level of severity of treatment as to avoid carbon deposition within the tubular reactor or producing an appreciable quantity of fixed hydrocarbon gases.
Bottoms from boot 3ft are recycled to the extent desired to heater-reactor 13 by means of lines 3S, 37 and E, J 11. Non-recycled bottoms may be withdrawn 'from the System through line 36 for further treatment or may be diluted with a lighter oil introduced through line 40 for use as a fuel. Any extraneous light oil or a light oil produced by the process of the invention may be used as the cutter oil introduced through line 40'.
In a preferred embodiment, overhead removed from tower 17 through line 31 is quenched by the addition thereto of oil from line 42. The quenched oil then passes through heat exchanger 44 where the temperature of the mixture is further reduced to below 800 F. preferably to about '700 to 750 F. The mixture is then transferred through line 45 to separator 46 from which liquid is removed through line 48 and gaseous material comprising hydrogen and vaporous hydrocarbons is removed through line 50. Optionally, a portion of the liquid material removed 'from separator 46 is recycled to heater-reactor 13 through lines 48, 37 and 11.
The gaseous overhead from separator 46 passes through line 50 to catalyst unit 52. Although the flow through catalyst unit 52 is depicted as down flow, the unit may also be operated up ow.
In catalyst unit 52, the mixture is contacted with a hydrogenation catalyst. Suitable catalysts comprise the oxides 'and/or suldes of metals such as cobalt, molybdenum, nickel, tungsten, chromium, iron, manganese, vanadium and mixtures thereof. The catalytic materials may be used alone or may be deposited on or mixed with a support such as alumina, magnesia, silica, zinc oxide and the like. Particularly suitable catalysts are nickel-tungsten sulfide, molybdenum oxide on alumina, a mixture of cobalt oxide and molybdenum oxide, generally referred to as cobalt molybdate on alumina, molybdenum oxide and nickel oxide on alumina, molybdenum oxide, nickel oxide and cobalt oxide on alumina, nickel sulde on alumina, cobalt sulde and nickel sulfide on alumina. Although these catalysts `are generally considered as hydrogenation catalysts, a considerable amount of conversion of the heavier hydrocarbons present into lighter boiling materials takes place in catalyst unit 52 probably due at least in part to the thermal treatment to which the feed stock has been subjected just prior to its introduction to the catalyst. The system pressure at which the reactant stream enters the catalyst unit 52 is usually satisfactory.
Effluent from catalyst unit 52 passes through line 53, cooler 54 and line 55 to high pressure separator 56 from which hydrogen is removed and recycled through lines 58, 21, and 11 to heater-reactor 13 and through lines 58 and 19 to heater 27. Liquid from high pressure separator 56 is transferred through line 59 to low pressure separator 60 from which light hydrocarbons are withdrawn through line 61 and heavier hydrocarbons are withdrawn through line 62. A portion of product in line 62 may optionally be recycled as quench through lines 64 and 42 or may be recycled to heater-reactor 13 with recycle tar through a line not shown or the entire liquid yield may be removed through line 65, and sent to fractionation or other treatment. Make-up hydrogen, as required, is introduced into the system through line 20.
The following example which is submitted for illustrative purposes only, shows how the feed, a vacuum residuum, is converted to greater yields of desired liquid products when increasing amounts of recycle bottoms are included with the fresh feed.
Gravity, API 5.4 Viscosity, SSF at 210 F 1237 Carbon residue, wt. percent 24.0 Sulfur, wt. percent 1.68 Nitrogen, wt. percent 1.33 Nickel, p.p.m. 183 Vanadium, p.p.m. 178
Run l Run 2 Run 3 Heater reactor outlet, F 900 900 900 Heater outlet, F 865 870 870 Heater reactor press., 1, 700 1, 700 l, 700 Tower pressure, p.s.i. 1, 625 1,600 1, 600 Catalyst unit inlet, F 775 770 775 Catalyst unit outlet, F 725 725 725 Catalyst unit inlet, p.s.i.g 1,560 1, 530 l, 540 Catalyst unit outlet, p.s.i.g 1,528 1, 500 l, .500 Oil feed, 1b./hr 770 668 360 Bottoms recycle, lb./hr 301 540 Hydrogen to heater-reactor s 14, 500 14,100 12,000 Hydrogen to heater s.c.f.h 34., O 33, 100 24, 000 Net Productsu 0.1-400" F., vol. percent basis fresh feed 4. 3 4. 9 10. 6 400-500 F., vol. percent basis fresh feed 3.1 3.1 38. 7 50G-670 F., vol. percent basis fresh feed 15. 4 l5. 2 670 F.-lvol. percent basis fresh feed- 15. 5 27. 4 51. 8 Bottoms, vol. percent basis fresh feed 63. 8: 52. 4 9. 9 Hydrogen consumption, s.c.f.b. fresh feed 307 296 1550 We claim:
1. A process for the hydroconversion of a hydrocarbon oil feed stock having a Conradson carbon residue of at least 1% by weight which comprises passing an intimate mixture of hydrogen and said oil through a tubular reaction zone under conditions of turbulent flow at a temperature between about 700 and 1000 F. and a pressure not less than about 500 and not greater than about 5,000 p.s.i.g., separating the reaction zone eluent into a gasiform portion comprising hydrogen and vaporous hydrocarbons and a liquid portion comprising unvaporized hydrocarbons, contacting the liquid portion with a hydrogen stream separately heated to a temperature between about 800 and 950 F. to strip dissolved and entrained ygaseous materials from said liquid portion and produce a tar-like viscous residue and introducing at least a portion of said tar-like viscous residue with fresh feed into the tubular reaction zone.
2. The process of claim 1 in which the intimate mixture contains at least 1,000 s.c.f. of hydrogen/bbl. of fresh feed.
3. The process of claim 1 in which the liquid portion is contacted with at least 5,000 s.c.f. of hydrogen per f bbl. of fresh feed.
4. The process of claim 1, in which a portion of the stripped lliquid is `subjected tto par-tial combusti-on tto produce hydrogen `for the hydroconversion of additional feed.
5. The process of claim 1 in which the gasiform portion is contacted with a hydrogenation catalyst at a temperature not greater than about 800 F.
6. The process of claim 5 in which the gasiform portion is cooled to a temperature below about 900 F. and the resulting condensate is removed from the stream prior to the introduction of the stream into the catalytic hydrogenation zone.
7. The process of claim 5 n which a portion of the liquid product recovered from the catalytic hydrogenation zone is added as quench to the gasiform portion.
8. The process of claim 6 in which a portion of the condensate is added as quench to the gasiform portion.
9. The process of claim 6 in which a portion of the condensate is recycled to the tubular reaction zone.
10. The process of claim 5 in which the turbulence level in the tubular reaction zone is at least 25.
References Cited bythe Examiner UNITED STATES PATENTS 2,853,433 9/1958 Keith 208--58 2,989,459 6/ 1961 Eastman et al. 208--107 2,989,461 6/1961 Eastman et al. 208--107 3,044,951 7/ 1962 Schliniger et al. 208-95 DELBERT E. GANTZ, Primary Examiner. ALPHONSO D. SULLIVAN, Examiner.