US 3228871 A
Description (OCR text may contain errors)
Jan. 11, 1966 w. G. SCHLINGER 87 TREATMENT OF HYDROCARBONS WITH HYDROCRACKING IN THE FIRST STAGE AND HYDROGENATION OF THE GASEOUS PRODUCTS Filed Aug. 7, 1962 United States Patent "ice TREATMENT OF HYDROCARBONS WITH HYDRO- CRACKING IN THE FIRST STAGE AND HYDRO- GENATHUN OF THE GASEOUS PRQDUCTS Warren G. Schlinger, Pasadena, Calif., assignor to Texaco Inc, New York, N.Y., a corporation of Delaware Filed Aug. 7, 1962, Ser. No. 215,380 9 Claims. (Cl. 208-58) 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 in the presence of hydrogen into lighter hydrocarbon oils of reduced carbon and/0r metal contents.
Several methods for the hydroconversion of heavy hydrocarbons into lighter hydrocarbons are known. In one particularly advantageous process for the conversion of hydrocarbon oils, the hydrocarbon oil, in intimate mixture with hydrogen, is passed through a tubular reaction zone at elevated temperature and pressure under conditions of highly turbulent flow. The eflluent from the tubular reaction zone is then introduced into a separation zone wherein the gaseous material is seperated 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 from the separated liquid material. The resulting gaseous stream comprising unreacted hydrogen and vaporous hydrocarbons is combined with the gaseous material removed .from the tubular reaction zone effluent and the combined stream is passed into contact with a hydrogenation catalyst. Such a process is disclosed in U.S. patent application Serial No. 33,582, filed June 2, 1960, US. Patent No. 3,089,843, of which I am co-inventor.
The proccess described above is characterized by the circulation of large volumes of hydrogen far in excess of the stoichiometric hydrogen requirements of the conversion. The total hydrogen rates range from 10,000 up to 100,000 s.c.f.b. (standard cubic feet per barrel) of feed, from 13,000 to 80,000 s.c.f.b. being preferred. It is also disclosed that from 10 to 60% of the hydrogen is introduced into the tubular reaction zone with the feed and that at least about 5,000 s.c.f.b. is introduced into the contacting zone. Although the known process is capable of converting heavy residual stocks into good yields of middle distillates a large hydrogen inventory is required and compression costs for the circulation of such large volumes of hydrogen are high.
It is an object of the present invention to convert hydrocarbon liquids into lighter hydrocarbon products. A further object of the present invention is to convert heavy hydrocarbon oils into good yields of middle distillates. A further object of the invention is to convert heavy hydrocarbon oils into middle distillates using less circulating hydrogen than is used in prior processes. A still further object of the present invention is to convert heavy hydrocarbon oils into good yields of middle distillates which are particlularly suitable feed stocks for thermal and catalytic cracking operations.
According to the present invention, the hydrocarbon liquid and hydrogen is passed in the form of an intimate mixture through a tubular reaction zone under conditions of turbuient flow, the efiluent from the tubular reaction zone is passed into a separation zone wherein the gas phase comprising hydrogen and vaporous hydrocarbons is separated from the liquid phase comprising unvaporized by- 3,228,871 Patented Jan. 11, 1966 drocarbons, at least a portion of the separated liquid phase is recycled to the tubular reaction zone, the recycled portion being maintained in the substantial absence of hydrogen until it is reintroduced into the tubular reaction zone and the gas phase separated from the tubular reaction zone efiiuent is passed into contact with a hydrogenation catalyst.
Any hydrocarbon liquid may be used as the fresh feed to the process of the present invention. In this respect, the term fresh feed identifies hydrocarbon material which is being introduced to the tubular reaction zone for the first time, as distinguished from recycle material. However the process has particular application in the treatment of hydrocarbon liquids containing residual components, metals and other tar and ash forming constituents. Hydrocarbon liquids for the treatment of which the process of the present invention is particularly adapted 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 and Arabian crude, naphtha, kerosene and heavy fractions of crude oils such as reduced or topped crude, deasphalted oil, vacuum residuum, and mixtures thereof and the like. Other materials which may be advantageously treated are coal oil, pitches, tars gilsonite, shale oil and tar sand oil.
The hydrogen employed in the process of the present invention may be substantially pure, e.g. -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 may be used after being treated with a water gas shift conversion catalyst and then scrubbed for CO 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 60 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 of turbulence, the feed or at least that portion which is liquid at the pre vailing conditions is in the form of fine mist-like droplets suspended in a gaseous medium comprising hydrogen. With the reaction mixture in this state, 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.
In the initial step of the process of the present invention, the. fresh feed, the recycle liquid and the hydrogen are introduced into a tubular reaction zone which is maintained at a temperature between about 700 and l000 F., preferably 800 to 925 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 1,000 and 2,000 p.s.i.g. Hydrogen is circulated in amounts in excess of the stoichiometric requirements of the process. Hydrogen rates of between about 1,000 and 50,000 s.c.f.b. of fresh feed are used, preferred rates being between 3,000 and 20,000 s.c.f.b.
The effluent from the tubular reaction zone which is in the form of a suspension of mist-like droplets of unvaporized oil suspended in a gaseous medium comprising hydrogen and vaporous hydrocarbons is passed to a separa- -tion zone which is maintained at substantially the same temperature and pressure as the tubular reaction zone. In the separation zone the unvaporized hydrocarbon is separated from the gaseous suspending medium. The separated gaseous phase is removed from the separation zone and is then passed into contact with a hydrogenation catalyst.
The separated liquid phase may be recycled directly to the tubular'reaction zone. Advantageously, however, the separated liquid is maintained for a period ranging from two seconds to two minutes at substantially the same conditions of temperature and pressure as the separation zone prior to being reintroduced with hydrogen and fresh feed into the tubular reaction zone.
While theoretically, it is possible to recycle all of the separated liquid to the tubular reaction zone, in practice it is desirable to withdraw a portion of the separated liquid from the system to prevent the buildup of metals and other undesirable constituents of the feed. Usually, a 5 to 20% draw off basis fresh feed is satisfactory. As applied here the term liquid used in connection with the separated liquid is intended to mean that the material is liquid at the prevailing conditions. In some instances depending on the nature of the feed and the severity of the reaction conditions, this material can be solid at room temperature.
The amount of separated liquid recycled to the tubular reaction zone will vary with the amount of liquid recovered from the separation zone. However, it is possible for the total hydrocarbon feed to the tubular reaction zone to contain as much as 70% by volume recycle liquid. It is advisable, during this hold-up period, to supply agitation to the separated liquid. This may be done by bubbling a gas through the liquid at a rate of about 100- 250 s.c.f. per barrel of liquid present preferably 100-200 s.c.f. Suitable gases are hydrogen, methane, nitrogen and the like.
It is a feature of the present invention that the increase in hydrocarbon input into the tubular reaction zone as a result of the recycling of the separated liquid does not require a corresponding increase in the hydrogen input.
For a better understanding of the invention, reference is made to the accompanying drawing which illustrates diagrammatically a preferred embodiment of the present invention.
Fresh feed introduced into the system through line 11 I and recycle liquid from line 19 are mixed with a large excess of hydrogen from line 21 and the mixture is introduced into heater-reactor 13 in which it is passed through a tubular reaction zone under conditions of highly turbulent flow. The hydrogen is mixed with the oil in an amount ranging from about 1,000 to 50,000 s.c.f.b. of fresh feed preferably from about 3,000 to about 20,000 s.c.f.b. Reaction temperatures within the tubular reaction zone are maintained between about 700 and 1,000 F. preferably between about 800 and 900 F. The tubular reaction zone outlet pressure is advantageously maintained within the range of about 1,000 and 2,000 p.s.i.g. although pressures ranging from as low as 500 to as high as 5,000 p.s.i.g. or higher may be employed. The
reactants pass through the tubular reaction zone as an intimate mixture in a state of high turbulence. The higher boiling hydrocarbons under the prevailing conditions of temperature, pressure, hydrogen to oil ratio and contact time are subjected to viscosity breaking with substantially immediate hydrogenation of the molecular fragments without further breakdown, thereby increasing the production of middle distillates boiling in the 400 to 700 F. range without the concomitant production of large volumes of fixed gases and low molecular weight hydrocarhens and coke. As the reaction proceeds the molecular fragments because of their low boiling characteristics are substantially immediately vaporized.
The hot mixture of hydrogen, vaporized hydrocarbons containing suspended mist-like droplets of unvaporized oil then passes through line 15 to the upper section of separator 17. This section serves as a disengaging zone to separate the gasiform materials comprising vaporous hydrocarbons and hydrogen from the unvaporized oil. The unvaporized oil collects'in the lower section of separator 17 where it is maintained in the liquid state in the substantial absence of hydrogen. Periodically or preferably continuously the separated liquid is withdrawn from separator 17 through line 18 and is recycled to tubular reaction zone 13 through lines 19 and 11, the separated liquid being maintained in the substantial absence of hydrogen until its introduction into heater-reactor 13. Gas for agitation is supplied to tower 17 through line 16. It is contemplated that the expression in the substantial absence of hydrogen includes the use of hydrogen for agitation purposes in amounts up to about 250 s.c.f. per barrel of separated liquid. Although the recycle liquid is reintroduced into the tubular reaction zone in addition to the fresh feed, it is a feature of the present invention that no additional hydrogen is used to compensate for the increased hydrocarbon input to the tubular reaction zone.
Overhead from separator 17 is quenched in line 31 by the addition thereto of oil from line 32. The quenched oil then passes through heat exchanger 33 where the temperature of the quenched mixture is further reduced to below 800 F., preferably to about 700-750 F. The mixture is then sent through line 34 to catalyst unit 36 which contains a hydrogenation catalyst. Suitable catalysts for use in hydrogenation unit 36 comprise the oxides and/or sulfides 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, or mixtures thereof. 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 sulfide on alumina, cobalt sulfide and nickel sulfide on alumina. While reactant flow through the catalyst unit is shown as down flow, it is also possible for the reactants to be passed upwardly through the catalyst unit.
Although these catalysts are generally considered to be hydrogenation catalysts, a considerable amount of conversion of the heavier hydrocarbons present into lighter boiling materials takes place in catalyst unit 36 probably due at least in part to the thermal treatment to which the stock has been subjected prior to its contact with the catalyst. Space velocity in catalyst unit 36 ranges from about 0.1 to 10 volumes of liquid fresh feed per volume of catalyst per hour. Preferably the space velocity will range be tween about 0.5 and 5.
Efliuent from catalyst unit 36 then passes through line 38, cooler 39 and line 40 to high pressure separator 41 from which a gas rich in hydrogen is removed and recycled to heater-reactor 13 through lines 21 and 11.
Liquid product is removed from high pressure separator 41 through line 43 through which it may be sent to a low pressure separator for the separation of normally gaseous hydrocarbons from normally liquid components which may then be sent to fractionation and if desired to further treatment. Optionally a portion of the liquid product may be removed from high pressure separator 41 through line 45 and recycled as quench through line 32 or used as cutter oil for material removed from the system through line 18. Make-up hydrogen may be added, as required, through line 14 to recycle hydrogen stream in line 21.
The following example is submitted to show the advantageous results obtainable by the process of the present invention while using much less hydrogen circulation than the process of the prior art.
Run 1, experimental data for which appear below, follows closely the sequence described in U.S. patent application Serial No. 33,582, filed June 2, 1960, now US. Patent No. 3,089,843. In run 1, the hydrogen-hydrocarbon mixture is passed through a tubular reaction zone at elevated temperature and pressure, the eflluent is introduced into a separation zone where the gaseous material is separated from the liquid material, the liquid material is contacted with a separately heated stream of hydrogen to eifect additional hydroconversion of the separated liquid and to remove vaporous hydrocarbons entrained and dissolved in the liquid, the hydrogen-vaporous hydrocarbon product stream from the contacting is combined with the separated gaseous material and the combined stream is passed into contact with a hydrogenation catalyst.
In run 2 the sequence follows closely that of the accompanying drawing. in run 2 the hydrogen-hydrocarbon mixture is passed through the tubular reaction zone and the effluent is separated into a vaporous stream and a liquid stream. By this separation unconverted hydrocarbon liquid is separated from the hydrogen and is kept out of contact with hydrogen either until it is reintroduced into the tubular reaction zone or immediately prior to its reintroduction into the tubular reaction zone. In run 1 liquid or tar recovered from the separation and contacting steps amounts to 23.6 volume percent basis fresh feed. In run 2 liquid from the separation step amounts to 44 volume percent basis fresh feed. To provide suitable data for comparison, at equilibrium conditions for every 100 barrels of fresh feed introduced into the heater-reactor in run 2, 37 barrels of tar are recycled to the tubular reaction zone. This gives a net tar yield, in both runs, of about 23% and shows that the recycle tar acts as fresh feed. Data for runs 1 and 2 are tabulated below.
Fresh feed boiling range:
IBP-400 F., vol. percent 1.5 400-600" F., vol. percent 15.7 600-1000 F., vol. percent 32.2 1000 F.+vol. percent 51.6
Run 1 Run 2 Pressure, p.s.i.g.:
Tubular reactor inlet.-. 1, 203 1, 205 Hydrogen heater inlet.. 190 Catalyst chamber outlet. 1, 028 1, 025 High pressure separator 984 980 Temperature, F.:
Tubular reactor outlet. 882 885 Hydrogen heater outlet- 898 Separation zone overhead 858 855 Catalyst chamber outlet 787 790 High pressure separator inlet 99 100 Recycle tar in feed, vol. percent 27. 0 Flow rates}, bafiisdfesh ieettl: t b 1 ct ecce oen o uuar rea or s.c l.b.. 3---? 8,450 8, 450 Recycle hydrogen to hydrogen heater s.c.Lb 4, 600 Make-up hydrogen s.c.f.b...- 500 500 Hydrogen content, vol. percent. 68. 1 68. 7 Yields basis oil charged:
(l -03w t. percent.-. 1.4 1.6 U i-oil, vol. percent. 79. 6 80. 4 Tar, vol. percent 23. 6 23. 3 IB P-400 I vol. percent. 14.0 14. 3 400500 F. v01. percentl6. 0 17. 1 500-700 F. vol. percent- 40. 3 41. 2 700 F.+vol. percent 29. 7 27. 4
1. A process for the hydroconversion of a heavy hydrocarbon oil having a Conradson carbon residue of at least 1% which comprises passing an intimate mixture of said oil and hydrogen through a tubular reaction zone under conditions of turbulent flow at a temperature between about 700 and 1000 F. and a pressure between about 500 and 5000 p.s.i.g., separating the effluent from said tubular reaction zone at prevailing conditions into a liquid portion and a gaseous portion, collecting the separated liquid and maintaining same at substantially the same conditions of temperature and pressure as the separation zone recycling at least a portion of said collected liquid portion to the tubular reaction zone, said recycled portion being maintained in the substantial absence of hydrogen from its separation until its reintroduction into said tubular reaction zone, cooling said gaseous portion, thereafter passing said gaseous portion into contact with a hydrogenation catalyst under hydrogenation conditions and recovering normally liquid hydrocarbons from the catalytic hydrogenation zone efliuent.
2. The process of claim 1 in which in the tubular reaction zone the temperature is between 800 and 925 F.
3. The process of claim 1 in which the intimate mixture contains between 1000 and 50,000 s.c.f. hydrogen per barrel of fresh feed.
4. The process of claim 1 in which the collected liquid is maintained at prevailing conditions of temperature and pressure for a period of time between 2 seconds and 2 minutes prior to its reintroduction into the tubular reaction zone.
5. The process of claim 4 in which the collected liquid is agitated during said period.
6. The process of claim 1 in which the hydrogenation catalyst is maintained at a temperature between about 700 and 800 F.
7. The process of claim 1 in which the hydrogenation catalyst comprises a compound of molybdenum supported on alumina.
8. The process of claim 1 in which the hydrogenation catalyst comprises cobalt molybdate on alumina.
9. The process of claim 1 in which the hydrogenation catalyst comprises the sulfides of nickel and tungsten.
References Cited by the Examiner UNITED STATES PATENTS 1,974,057 9/1934 Steffen et al. 20854 2,989,461 6/1961 Eastman et al. 208107 3,008,895 11/1961 Hansford et al. 208-l12 3,119,765 l/1964 Corneil et al. 208-59 3,148,135 9/1964 Schlinger et al. 208-58 DELBERT E. GANTZ, Primary Examiner.
ALPHONSO D. SULLIVAN, Examiner.