|Publication number||US3193595 A|
|Publication date||Jul 6, 1965|
|Filing date||Jan 31, 1962|
|Priority date||Jan 31, 1962|
|Publication number||US 3193595 A, US 3193595A, US-A-3193595, US3193595 A, US3193595A|
|Inventors||Joseph R Kenton, Richard E Pierson|
|Original Assignee||Standard Oil Co|
|Export Citation||BiBTeX, EndNote, RefMan|
|Patent Citations (3), Referenced by (10), Classifications (13)|
|External Links: USPTO, USPTO Assignment, Espacenet|
y 1965 J. R. KENTON ETAL 3,193,595
HYDROCARBON CONVERSION 2 Sheets-Sheet 1 Filed Jan. 31, 1962 le Toluene INVENTORS: Joseph R. Kenton Ric/lard E. Pierson MQC-QQMM ATTORNEY Air h dro en leed 7 Hydrocarbon July 6, 1965 .1. R. KENTON IETAL HYDROCARBON CONVERS ION 2 Sheets-Sheet 2 Filed Jan. 31, 1962 WSQMMXW Nu mxlwm kvnSQ QW QQQNQ mww United States Patent M 3,193,595 HYDROCARBON CONVERSION Joseph R. Kenton, Tulsa, Okla, and Richard E. Pierson,
Country Club Hills, lll.,, assigors to Standard Oil Company, Chicago, EL, a corporation of Indiana Filed Jan. 31, 1962, Ser. No. 170,169 12 Claims. (Cl. 260-672) This invention relates to the conversion of aromatic hydrocarbons, especially hydrocarbons containing a major proportion of alkyl aromatic hydrocarbons. More specifically, this invention pertains to the conversion of aromatic hydrocarbon fractions containing a major proportion of alkyl aromatic hydrocarbons into useful aromatic hydrocarbon product fractions including aromatic hydrocarbon product fractions enriched with respect to unsubstituted aromatic monocyclic hydrocarbons and to unsubstituted aromatic polycyclic hydrocarbons. This invention also pertains to a novel integrated system for said aromatic hydrocarbon conversion employing thermal noncatalytic hydrodealkylation of the alkyl aromatic hydrocarbon fraction and separation of the dealkylated feed into high equality unsubstituted aromatic hydrocarbon products.
Early investigators of the thermal stability of alkyl aromatics, such as toluene, cymenes, xylenes, naphthalenes and the like reported that these alkyl aromatics could be converted to dealkylated products by passing substantially pure alkyl aromatic hydrocarbon through a heated tube with or without such inert packing as carbon, glass beads, quartz chips, etc. Later researchers reported that by employing a mixture of hydrogen and alkyl aromatic hydrocarbon the thermal dealkylation appeared to be favored. However, none of these investigations was developed beyond the laboratory processes employed to study the thermal stability of alkyl aromatic hydrocarbons.
Later attention was directed to the effect of various catalytic materials on the thermal stability of alkyl aromatic hydrocarbons in the presence and in the absence of hydrogen. Even in the presence of catalytic materials, hydrogen favored the dealkylation. It was found from these early studies of the thermal stability of alkyl aromatics that dealkylation occurred at lower temperatures in the presence of a catalyst than in the absence of a catalyst. Also the early researchers reported that the alkyl aromatics were converted to coke in the presence or absence of catalyst as well as undergoing dealkylation.
Many proposals have been offered to improve the thermal catalytic and non-catalytic dealkylation of alkyl aromatic hydrocarbons in the presence of hydrogen to improve the yield of unsubstituted aromatic hydrocarbon and to reduce coke formation.
The improved process of this invention relates to the thermal non-catalytic hydrodemethylation of methyl substituted aromatics containing 1 to 2 aromatic rings, preferably toluene and methyl naphthalenes, to produce a reaction efiiuent containing benzene and naphthalene, respectively, and the separation from the reaction effluent of a gaseous mixture of hydrogen and paraflins, dealkylated aromatic benzene or naphthalene, methyl aromatics for recycle and high boiling aromatics. The improvement in the process comprises heating the charge stock; that is, the mixture of hydrogen and methyl aromatic feedstock, by indirect heat exchange with the reaction effluent to a temperature in the range of 300 to 400 F. below the dealkylation reaction temperature and thereafter separately heating hydrogen and feedstock with the partially cooled reactor effluent, preferably heating the hydrogen before heating the feedstock, and preferably employing portions of the cooled reaction efiluent after heating the charge stock to supply heat to the reboilers of 3,l93,55 Patented July 6, 1955 the distillations where the gaseous mixture of hydrogen and parafiins are stripped from the reaction effluent and where at least the unsubstituted aromatic product; benzene or naphthalene, is fractionated from the remainder of the reactor eflluent; thereafter heating the preheated charge stock to a temperature in the range of 1100 to 1340 F. in a combination of parallel flow paths, preferably free flow paths, by radiant heat, maintaining the charge stock at a temperature in the range of 11 to 1340 F. in series free flow path, preferably at substantially isothermal conditions within a maximum deviation of i20 F. from the dealkylation temperature, to convert at least of the methyl aromatics, preferably toluene and methyl naphthalenes, in the feedstock to unsubstituted aromatics: benzene or naphthalene.
In the foregoing improved process the thermal hydrodemthylation of toluene to benzene and methyl naphthalenes to naphthalene are carried out .at temperatures desirably in the range of 1100 F. to 1250 F., and preferably 1150 to 1300 F., at hydrogen pressures suitably in the range of from 450to 1000, desirably 600 to 900 and preferably 700 to 800 pounds per square inch gauge (p.s.i.g.) at ratios of hydrogen to methyl aromatic feedstock suitably 5000 and above, desirably 5000 to 11,000 and preferably 6000 to 9000 standard cubic feet of hydrogen per barrel of methyl aromatic feedstock. The flow rate of the charge stock in the reaction zone should be at a volume hourly space velocity rate expressed as volume (V of methyl aromatic feedstock per hour per volume of reactor (V Suitable volume hourly space Velocities should be above 0.5 Vg/lJL/V for below this value excessive coking results from increased residence time in the reaction zone. Desirable volume hourly space velocities are in the range of 0.5 to 5.0 with the higher space velocities being associated with the lower tempera tures and the lower space velocities being associated with the higher temperatures of hydrodemethylation hereinbefore set forth for the process of this invention. Volume hourly space velocities preferred for the process of this invention are in the range of 0.5 to 3.0. High space velocities, those above 3 to 5, in general, reduce the rate of conversion by hydrodemethylation but can be used where recycle of methyl aromatics in greater volume is practicable. In the preferred range, high conversions above 50 mole percent and up to mole percent and higher, with a single pass operation can be obtained. In the reaction zone catalytic materials supported or unsupported are, of course, completely absent. However, substantially non-catalytic packing such as quartz or high silica (96% silica) chips, beads, low pressure drop glass fiber packing, stainless steel chips, steel wool and the like can be used, if desired. Such packing materials are rather poor heat transfer materials; they tend to restrict flow of reactants in the reactor tube and provide coking sites. The use of unpacked tubes in the reactor-furnace hereinafter described are preferred. Demethylation temperatures below about 1075 F. result in demethylation conversions too low to be commercially acceptable. No appreciable reaction takes place at 1050 F. Demethylation temperatures in excess of 1340 F. enhance cracking and other competing reactions as well as enhancing rapid coke formation. The improved process, as hereinbefore described, avoids the formation of coke, especially when conducted in the temperature range of 1100 F. and 1300 F.
One pass systems, that is, where the alkyl aromatic containing feedstock is passed once through the thermal catalytic dealkylation step, although causing dealkylation to occur, could not produce high yields of the dealkylated product. For example, in our laboratories a monomethyl naphthalene fraction (containing both alpha and beta methyl naphthalene), when passed over a chromia-alumina catalyst at 950 to 1150 F. at a space velocity of pound per pound of catalyst per hour with hydrogen in the ratio of 3800 to 8800 cubic feet (volume measured at S.T.P.) per barrel of monomethyl naphthalene and a hydrogen pressure of 500 to 1000 p.s.i.g. produced naphthalene yields up to 30 Weight percent based on the aromatic feedstock. Ourexperiences showed that by increasing the dealkylation temperature, coking increased; increasing the hydrogen ratio to the hydrocarbon did not reduce the coking at temperatures above 1150 F.; and increasing the dealkylation pressures did not reduce coking or enhance the yield of naphthalene in such a catalytic thermal dealkylation process. I
However, it was found in our laboratories that substantially enhanced hydrodealkylation could beobtained I 4 V loss to gas and handling of 26.9%. The handling loss does not exceed about 3 to 5%.
When commercial hydrogen is charged with the mixture of mono-, diand trimethylnaphthalene feedstock hereinbefore employed, the off gas composition separated from the aromatic condensate from the dealkylat the higher temperatures required by non-catalytic thermal hydrodealkylation provided the residence time at the higher temperatures Was controlled. It'was found that advantageous control could be provided by passing hydrogen and alkyl aromatics through a tube. packed with quartz or Vycor chips, stainless steel chips or steel wool at temperatures in the range of 1200 to 1400 F., hy-
drogen pressures in the range of 400 to 1000 p.s.i.g. and
the first pass without coke formation. Recycling of the unconverted methyl naphthalenes under the same conditions produced more naphthalene. Recycle of the un-' converted mcthylnaphthalenes from the second pass to a third pass produced still more naphthalene. After a a total of live passes under the same conditions there still was no coking evident even though 99.3% of the original charge had been converted. The total yield of naph ation ty-pically'contains on a mole percent basis:
Compdn'ent- Mole percent Hydrogen 85 to 87 Methane 11 to 13 Carbon monoxide 0.3 to 0.4
. Ethylene 0 to 0.1 Ethane 1.2 to 1.4 Propane 0 to 0.2 Isobutane 0 to 0.1
Ethane is probably present from either deethylation or by combination of methyl radicals. e a Commercial hydrogen need not be employed in the hydrodealkylation reaction. The above off gas contains .sufiicient hydrogen to be recycled without purification. Also hydrogen streamsreadily available in refineries such as thehydrogen streams from catalytic reformer processes are excellent sources of hydrogen for hydrodealkylation. Such hydrogen containing reformer gases contain 80 to 85 mole percent hydrogen'with the remainder being low molecular Weight hydrocarbons, mainly methane. I
Still another means for controlling the retention time at 1100 to '1340-F. according to this invention is to provide rapidfiow and heat exchange between a higher temperature heating means, temperatures higher than the dealkylati-on temperature, and then hold at a lower flow rate the mixture of hydrogen and alkyl aromatic hydrocarbon at the dealkylation temperature. While packed tubes can be advantageously employed in the laboratory, pilot plant and smallscale commercial operations, theuse of free or unpacked tubes is preferred for large scale 1 commercial operation- The above concept of the use of thalene was 94% of theoretical based on the methyl naphthalenes in the feedstock mixture. The resulting hydrocarbon mixture produced when fractionated results in a pro-naphthalene fraction, materials boiling below naphthalene, a 79 C. naphthalene and a bottoms fraction. The pre-naphthalene fraction contains on a weight basis:
1 Percent Benzene 12.5 Toluene 20.0
Xylene 25.0 Other aromatics 42.5
and is found by 'sulfonation to contain 98.6 volume percent aromatics. The research F-1 octane number of the pre-naphthalene fraction is 105.6, makingit an excellent component for premium gasoline blending.
Using the same feedstock described above containing mono-, diand trimethyl n-aphthalenes but using glass chips (96% silica) and chips of303 stainless steel as packing, a hydrogen pressure of 450 p.s.i.g., a hydrogen to feedstock ratio of 6300 cubic feet per barrel and a de alkylation temperature of 1328" F., the weight yield' of naphthalene on feedstock charged is 39 to 40% over a space velocity of 0.5m 4.0 on a one pass basis. No coke formation is apparent. Under substantially the same conditions but at about 1300 F. dealkylation temperature the naphthalene yield is about 30 weight percent based on the .feedstock charged. The same feedstock with hydrogen at 6300 cubic feet/bbl. and a hydrogen pressure of 800 ,p.s.i.g., passed through a tube packed with the high silica glass chips (96% silica) at a dealkylation temperature of 1250 to 1300" F. results in weight yields based on the feedstock of 5.3% pre-naphthalene, 51.5% naphthalene, 16.3% post-naphthalene fraction and free tubes to first provide a minimum of time of exposure to high skin temperatures and a lower flow rate at skin temperatures at or about dealkylation temperature is a unique feature of the integrated system of this 1 invention for by'the utilization of this concept coke formationin the heating and reaction tubes can be avoided.
"Coking does not become a problem when inert packing, especially stainless steel chips, steel wool, quartz chips, high silica (96% silica) chips or beads, are employed until dealkylation temperatures in excess of 1340 1F. are employed. At temperatures of from 1200 to 1340 F.' schedule 80 stainless steel tubes packed with inert packing can be employed to provide the reaction zone for hydrodealkylation without formation of coke on the packing or on the inner tube surfaces. The use of full 'fiowtubcs, unpacked tubes, when employed according to the unique arrangement hereinafter, described, permits operation of the hydrodealkylation reaction at high conversions of methyl aromatics to unsubstituted aromatics, especially the conversion of a methyl naphthalenes to naphthalene and toluene to benzene, as high as without recycle and with substantially no coke formation athydrodealkylation temperatures of up to 1340 F.
There are a wide variety of possible unpacked tube reactor designs. Suitable unpacked tube reactors useful for the process of this invention are those having a ratio of length (L) to internal diameter (D), both in inches, in the range of from 15 to 800. Preferably the tube reactor L/D ratio is above 18. Tube reactor L/ D ratios lower than-about 15 tend to produce back mixing, that is reaction products tend to circulate back through the reactor tube, and reduce initial rates of reaction in the initial portion of the tube where the reactants are introduced thus resulting in low total conversions. It has been found that in unpacked tube reactors having L /Dratios of 15 and above there'is no apparent back mixing. It has V also been found that there is no substantial diiference in conversion between the use of a tube reactor (unpacked) having L/D ratio of about 18 and the use of a tube reactor (unpacked) having L/D ratio of about 560. The tube reactor for unpacked operation can be, for example, of schedule 80 type 316 or 317 stainless steel tubing of from A to 5 inches internal diameter and of a suitable length to provide the required volume for the designed through-put at the volume having space velocities hereinbefore disclosed. A particular unpacked tube reactor will be hereinafter described.
Theoretically one mole of hydrogen is consumed per mole equivalent of alkyl group split off. As the art well recognizes it is advantageous to employ an excess of hydrogen for the hydrodealkylation processes up to moles per mole equivalent of alkyl group to be split 05. Generally, the amount of hydrogen employed is in the range of 1 to 10 moles per mole of alkyl aromatic in the feedstock, preferably in the range of 5 to 7 moles per mole of alkyl aromatic in the feedstock. Expressed in terms of commercial flow rates the hydrogen or gas containing at least 80% molecular hydrogen is employed in the range of 3500 to 9000 cubic feet (measured at standard temperature and pressure and, hence, referred to as standard cubic feet: s.c.f.) per barrel of alkyl aromatic containing feedstock.
This invention is directed to the utilization of thermal non-catalytic hydrodealkylation of alkyl aromatic hydrocarbons, especially the thermal non-catalytic hydrodemethylation of toluene and methyl naphthalenes to benzene and naphthalene, respectively, in an integrated process,
which, in a special case wherein the highly refractory bottoms fraction obtained from catalytic reformer processes is used as a source of methyl naphthalene feedstock, includes pretreatment of the feedstock source to eliminate some non-aromatic light ends and undesirable heavy ends in such catalytic reformer bottoms fractions. One point of uniqueness in the integrated process is the removal of heat from the dealkylation reaction mixture as the sources of heat for fractionator reboilers used in the separation of the aromatic components of the mixture resulting from the hydrodealkylation reaction and/ or as a source of heat for preheating the charge, hydrogen plus alkyl aromatic hydrocarbons, before bringing the charge to reaction temperature.
Another point of uniqueness is the manner in which the charge is brought from preheat temperature to reaction temperature. This is accomplished after and is dependent upon the preheating of the charging mixture. The hydrogen source is compressed to a pressure below that employed in the hydrodealkylation. The alkyl aromatic feedstock is pumped to about the pressure of the hydrogen source. The dealkylation reaction efiiuent is employed to heat the dealkylation charge stock in indirect heat exchange relationship by countercurrent flow; that is, the dealkylation reactor effluent substantially at dealkylation temperature heats the mixture of feedstock and hydrogen to a temperature within 300 to 400 F. of dealkylation temperature. This is important for the reason that the temperature gradient required to bring the charge stock to dealkylation temperature need not be so severe as to result in tube Wall temperatures at or about the temperatures, in excess of 1340 F., to cause coking. After heat exchange between the reactor effluent and charge stock, the cooler reactor eflluent is employed next to heat the pressurized hydrogen source and then to heat the alkyl aromatic feedstock. Another application of the same advantageous counter-current preheating of the charge stock and the components thereof involves also extracting heat from the reactor effluent, also before gas separation, to supply heat to the reboilers of the fractionators required to remove from the reactor efiiuent materials boiling below the dealkylated alkyl aromatic, benzene or naphthalene and to separate for recovery benzene or naphthalene. The use of the heat in the reacmic.
tor efliuent in the manners described greatly reduces utility requirements for a commercial plant and simplifies the plant system into a compact, unitary, integrated, substantially independent unit. Such a compact unitary system can be readily integrated, for example, with a catalytic reformer system .which supplies both the hydrogen and the feedstock methyl naphthalenes for the production of naphthalene as well as toluene for the production of benzene.
The foregoing heat exchange system, although exceptionally advantageous in a system employing thermal non-catalytic hydrodealkylation, can also be advantageously employed in a system for thermal catalytic hydrodealkylation of alkyl aromatics. In both cases not only is there eflicient utilization of the heat content of the reactor efiiuent, but there is also provided the benefit hereinbefore described with respect to preventing high skin temperatures enhancing coking conditions.
As hereinbefore indicated another unique feature of the integrated system of this invention making successful the integrated process employed therein is the furnace-reactor combination upon which is dependent the non-coking thermal non-catalytic dealkylation process. The furnace: reactor comprises a heating zone and a reaction zone. In the heating zone the charge stock flows rapidly through unpacked tubes in the radiant heating portion of a furnace heated by burners supplied with fuel such as a hydrocarbon fuel, or, preferably, a portion of the gas stripped from the reactor efiluent. The use of the gas removed from the reactor etfiuent as fuel is preferred, since its use also provides for compactness and independence of the unitary integrated system. Rapid flow through the heating zone in the furnace-reactor can be accomplished in many ways; however, it is preferred to pass the charge stock to be heated to reaction temperature through a plurality of unpacked tubes in parallel flow. In this man ner there is not only provided high contact area for transfer of heat but there is also minimized the exposure to heat up conditions further reducing the possibility for coke formation. The short residence time in the heating zone is made possible not only by the high flow rate of the charge stock therethrough but also by the preheating of the charge stock to Within 300 to 400 F. of dealkylation temperature. Such a final heating step can also be used to advantage where thermal dealkylation is carried out in the presence of a catalyst in the reaction zone. However, in the system of this invention it is preferred to employ in the reaction zone unpacked tubes or coils. The lack of packing materials (non-catalytic as well as catalytic) particles, chips, pellets, etc. not only provides a free unencumbered flow path but eliminates the possibility of sites to hold components of charge stock, especially the higher molecular weight components, and cause the trapped hydrocarbons to be carbonized. As is well known to the art, the hydrodealkylation of alkyl aromatics is exother- Hot spots can and to occur in or on a catalyst bed. By employing no catalyst in the reaction zone, another coking potential is eliminated. More important, the charge stock can be retained in the reaction zone at the volume hourly space velocities of 0.5 to 5 until to or higher conversions of the alkyl aromatic are accomplished even though the non-catalytic dealkylation temperature is somewhat higher, to 200 F. higher, than catalytic dealkylation temperatures.
The charge stock is permitted to flow more slowly in the unpacked tubes of the reaction zone than in the heating zone. To permit this slower flow there is provided in the reaction zone tubes or coils in series relationship so that the entire charge stock flows through all of the reactor free space rather than only a small fraction thereof as in the tubes or coils in the heating zone. To remove heat generated by the exothermic dealkylation reaction, air is mixed With the hot flue gases from the radiant zone of the furnace to provide an average temperature in the reaction zone Well below that 7 V at which coking may begin. Advantageously, the reactor tubes or coils in series are placed in the furnace convection section and air is added to the hot gases from the radiant zone ahead of the convection zone. Cocurrent flow of charge stock and fiue gas may be most desirable for maintainingisothermal conditions of minimum temperature deviation, say i20 F. However, a combination of cocurrent and countercurrent flow may be also employed with additional air tempering of flue gas for the countercurrent section of flow, if necessary. Cocurrent flow of tempered flue gas and charge stock in the reaction 'zone is preferred because the rate of reaction is highest at the reactor inlet and more heat of reaction needs to be removed. The reaction rate decreases and heat of reaction decreases toward the reactor outlet.
Another method for maintaining isothermal conditions in the reaction zone with a minimum of temperature deviation is to inject a portion, to 20%, of the methyl aromatic feedstock before admixing with hydrogen into the reaction zone coil downstream from the inlet to the reaction zone. The higher portion of the feedstock may be injected closer to the inlet and the lower portion of the feedstock may be injected farther from the inlet. The portion of the feedstock so injected into the reaction zone can be at the temperature conditions prevailing at the feed pump before contact with the reactor efiluent or at the temperature conditions following partialpre- 8 catalytic reformer gas is discharged through gas transfer line-18 to heater 19 and heated to 820 F. Catalytic reformer gas is employed in the ratio. of 8340 s.c.f. per barrel of toluene. The hot compressed catalytic reformer gas leaves heater 19 by gas transfer line 20 discharging into conduit 15 and mixing with. the heated toluene to form the charge stock. The charge stock at 550 F., the temperature resulting from mixing the heated catalytic reformer gas and toluene, passesthrough heat exchanger 21 where the'charge stock is heated to 900 F. and 77s p.s.i.g. The heated charge stock-flows through charging conduit 22 to furnace-reactor 23 which contains heating zone 24 and reactor zone 25. Heating zone 24 contains coil 26 in the radiant heating section of said zone. Heating zone 24is heated by burning fuel supplied burner 27 with air as shown. To supply the 19.6MM B.t.u./hr.
there is employed the necessary number of burners to burn, as the preferred fuel, a portion of the gas stripped from the reactor effluent. This gas is taken from gas line. The hot charge stock is heated in coil 26 to 1300 F.
' and passes into reactor coil 28 in reactor zone which heatingby the reactor effluent. I In either case the injected portion of the feedstock need not be admixed with hydrogen before being added to the portion of the reaction zone but the injected portion must, of course, be brought up to the pressure in the reaction zone by an auxiliary pump. It is also advantageous to add a portion of the feedstock not preheated to reaction temperature to any portion of the reaction zone wherein the charge stock is flowing countercurrent to the flue .gas. Where all the flow in the reaction zone is cocurrent with the flue gas, 15% of the feedstock may be injected at a point 15 to 20% of the length ofthe reaction zone downstream from the inlet of the charge stock heated to reaction temperature in the heating zone of the furnace-reactor.
The reactor efiluent after being cooled by indirect heat exchange in the fractionator reboilers and/or with the charge stock, hydrogen and feedstock, according to this invention, is then processed to separate the gases for fuel, stripping and recycle if desired, to remove hydrocarbons boiling below the dealkylated product benzene or naphthalene, to recover benzene or naphthaleneproduct from low boilers and unconverted alkyl aromatics for recycle. FIGURE 1 of the accompanying drawings is a schematic fiow sheet illustration of one system for utilizing the improved process of this invention and illustrates one embodiment of the improved integrated system of this,
methane making up most ofthe remaining 15 mole percent in the following manner. Toluene is taken from storage (not shown) by conduit 10 by pump 11 at 4635 gallons per hour and discharged at 795 p.s.i.g. and 100 F. into conduit 12 which feeds heat exchanger 13 through which flows in indirect heat exchange relationship reactor efliuent at 770 F. and 730 p.s.i.g. Toluene leaves heat exchanger 13 through conduit 14' and discharges into conduit 15. Hydrogen feed (catalytic reformer gas) is collected at 100 to 200 p.s.i.g. and is taken through conduit 16 by compressor 17 and compressed to 795 p.s.i.g. and 300 F. The compressed is the convection portion of the furnace.' Air is admixed with the flue gas from heating zone 24 by nozzles 29. The flow of air is adjusted to maintain an average charge stock temperature of 1300 F. in reactor coil 28 with a minimum deviation of 120 F. in reactor coil 28. Tempering air need not be added at the bottom of reactor zone 25 as shownbut can be added at the discharge connection between heating zone 24 and reactor zone 25 or into the lower .portion of the side of reactor zone 25 not common to and separating it from heating zone 24.
The reaction eflluent leaves reactor coil 28 by discharge line 30 at 1300 F. and 750 p.s.i.g. Reactor eflluent discharge line 30' conducts the hot reactor effluent to heat exchanger 21 where said efiluentis cooled to 980 F. and leaves at 740 p.s.i.g.1 after preheating the charge stock through conduit 31 and then heats catalytic reformer gas in heater 19 discharging therefrom at 770 F. and 730 p.s.i.g. to heat exchanger 13 to heat the toluene. The react-or effluent after heating the toluene is at 480 F. and 720 p.s.i.g. and still contains unused hydrogen, methane from demethylation and from the reformer gas source of hydrogen feed. The hydrogen content of reactor effluent is recovered.
In the schematic flow sheet of FIGURE 1 there are 7 shown three fractionating towers each with an external reboiler system. The total heat requirement for these reboiler systems is about 25 to 255x10 B.t.u./hr. for
the conversion of 4635 gallons toluene per hour to benzene. Before the reactor eflluent at 720 p.s.i.g. and 480 F. is'processed to remove hydrogen and parafl'lns, it is desirable for this scale of operation to remove about 16.5 X 10 B.t.u./ hr. from the reactor efiluent. This heat can be removed by air cooled heat exchanger 34, or by supplying some of the heat required by the reboilers directly in indirect heat exchange relationship in certain of the reboilers or by indirect heat exchange with a heating fluid which can then be further heated by burning a portion ofthe gases stripped from the reactor efiiuent. The latter means for additional heat removal will be hereinafter described in detail with respect to FIGURE 2.
Returning to the system illustrated in FIGURE 1, the reactor eflluent at 720 p.s.i.g. and 480 'F. is cooled by air passing through heat exchanger 34 and the efliuent, cooled to 200 F. and at 710 p.s.i.g., is charged to liquid-gas separator 36 bytransfer line 35. The separated gases leave the top of liquid-gas separator 35 by gas transfer line 37 and pass through cooler 38 cooled by water to F. and flows to separator 39 through conduit 40. Theliquids in liquid-gas separator 36 are under 710 p.s.i.g. pressure and are transferred to stripping tower 43 through conduits 41, 42. Stripping tower 43 is operated with a vapor pressure of 35 p.s.i.g. and F. at the top and'a bottoms temperature of 275 F.
The gaseous mixture leavingstripping tower 43 by v 9 vapor line 44 passes through cooler 45 from which condensate and uncondensed gases leave by conduit 46 and flow to liquid-gas separator 47. The liquid condensate is withdrawn from the bottom of liquid-gas separator 47 by pump 49 through conduit 48 and discharged through conduit 50 to the upper portion of stripping tower 43 as reflux. Heat is supplied to the liquid hydrocarbons in stripping tower 43 by reboiler 50 through which flows liquid hydrocarbons withdrawn from the lower portion of stripping tower 43 by conduit 51. The heated hydrocarbon mixture returns to stripping tower 43 via conduit 52. Heat may be supplied to reboiler 50, about 1.9 to 2.0 1O B.t.u./hr., by any suitable material flowing through reboiler 50 in indirect heat exchange with the hydrocarbon mixture withdrawn from stripping tower 43. Sufficient reaction effluent from conduit 33, for example, can be used as a source of heat according to a feature of this invention.
Uncondensed gases are removed from the topof liquidgas separator 47 by vapor transfer line 53, are compressed to 710 p.s.i.g. and discharged by conduit 54 to gas transfer line 37. Liquids collected in separator 39 are charged to stripping tower 43 via conduits 55 and 42. The gases removed from separator 39 via transfer line 56 contain mainly hydrogen and methane (at least 99 mole percent total) with the remainder being higher paraffins. A portion (about 6 to 6.5%) can be charged to furnace-reactor 23 as fuel from transfer line 56a and the remainder can be used for fuel or a source of hydrogen. When the hydrogen feed contains 85 mole percent hydrogen, the gas from separator 39 contains about 59 mole percent hydrogen when toluene is being demethylated.
From the bottom of stripping tower 43 a mixture containing benzene, toluene and higher boiling hydrocarbons is withdrawn by pump 60 by conduit 59 and charged to fractionator 62 operated with a vapor pressure of 50 p.s.i.g. at 280 F. and a bottoms temperature of 340 F. Under these conditions benzene product, about 3530 gallons per hour, is recovered by transferring the vapors (50 p.s.i.g. at 280 F.) by transfer line 63 through cooler 64cooled by air and thence through conduit 65 to condensate receiver 66. A portion of the condensate is recycled to fractionator 62 by pump 68 which withdraws condensate by conduit 67 and discharges condensate through conduit 69. Product benzene is removed from conduit 69 by product line 70, is cooled by product cooler 71 and sent to storage or benzene stripping by benzene transfer line 72.
Heat is supplied fractionator 62 by reboiler 75 which heats a toluene-heavy hydrocarbon mixture removed from fractionator 62 by conduit 74 and returned by con duit 7d. To recover the above amount of benzene, about 16.9 B.t.u./hr. heat is required by reboiler 75. This may be supplied by any heat transfer medium which will heat by indirect heat exchange in reboiler 75 the hydrocarbons fiowing therethrough to 340 F.
Bottoms from fractionator 62 are charged as feed by conduit '77 to toluene fractionator 80 operated with a vapor pressure of p.s.i.g. at 250 F. at the top and a bottom temperature of 340 F. Toluene is taken as an overhead fraction by passing vapors from the top of toluene fractionator 80 by transfer line 8]. to air cooled condenser 82, where a toluene condensate (140 F.) is obtained and flows via conduit 83 to receiver 84. The toluene condensate is recycled to toluene fractionator 80 in part as reflux by withdrawing condensate from receiver 84 by conduit 85 and pump 86 which discharges into conduit 87. The remainder of the toluene fraction is recycled to hydrocarbon feed conduit by conduit 88. Hydrocarbons and tars higher boiling than toluene are withdrawn from toluene fractionator 80 by heavy bottoms conduit 89 and pump 90 which discharges the heavy bottoms through conduit 91 through heavy bottoms cooler 92 and then to storage.
Heat is supplied to toluene fractionator by reboiler 94 through which flows liquid hydrocarbons taken from the lower port-ion of toluene fractionator 80 by conduit 93 and returned to said fractionator by conduit 95.
The process illustrated by FIG. 2 is especially adapted for the non-catalytic hydrodemethylation of methyl naphthalenes readily available from petroleum refinery streams and more particularly to the demethylation of bottoms fraction obtained by removing gasoline-boiling-range components from catalytically reformed naphtha or catalytic reformate. Briefly, catalytic reforming is carried out with a naphtha charge stock having a boiling range range of from 100 to 400500 F. This charge stock is reformed at 880 to 1000 F. and 50 to 800 p.s.i.g. in the presence of hydrogen and a catalyst having dehydrocycliza-tion activity.
After reforming, the mixture of'reformed hydrocarbons and recycle-hydrogen-containing gas is ordinarily cooled to separate liquefiable hydrocarbons as a gross liquid product from the hydrogen-containing gas for recycle. This liquid product contains a mixture of volatile hydrocarbon light ends, hydrocarbons in the gasoline boiling range; i.e., about 100 initial boiling point to about 350-420 F. final boiling point in the ASTM distillation, as well as a higher boiling material. This higher boiling material, the amount and composition of which depending upon the naphtha charging stock distillation range and the severity of reforming, is or contains the catalytic reformate bottoms fraction or fractions which are subsequently demethylated in accordance with the present invention. This bottoms fraction is variously termed polymer, post-gasoline, rerun bottoms, reformate botts or bottoms, etc., and is composed almost entirely (98 100%) of aromatic compounds, predominantly condensed ring aromatics. It is not presently known, nor is it important, precisely how this high boiling material originates in the reforming process.
The reformate bottoms fraction may be separated from gasoline boiling range hydrocarbons (either before or after stabilization of such hydrocarbons to remove light ends) by conventional fractional distillation in mul-ti-tray distilling or rerun towers. The operation of these towers may be controlled to provide any desirable endpoint in the gasoline boiling range hydrocarbons which are taken overhead.
The reformate bottoms as it is taken from the rerun tower contains hydrocarbons which boil within the range suitable for providing the 'demethylation feedstock herein. Typical reformate bottoms, a mixture of aromatics, is one having a gravity of 9.5 degrees API and the following ASTM distillation characteristics.
ASTM distillation: Degrees Fahrenheit Initial boiling point 460 10 percent point 480 30 percent point 500 50 percent point 508 70 percent point 523 percent point 588 Final boiling point 700+ Also useful in the preparation of naphthalene by thermal hydrodemet-hylation is an aromatic portion, 40 to 50 volume percent, of light catalytic cycle oil boiling over the range of 400 to 550 F. This mixture contains aromatics, parafiins, naphthalenes and sulfur compounds. By extracting light catalytic cycle oil (LCCO) with sulfur dioxide, a parafiin-naphthene raflinate and an aromatic-sulfur com-pounds extract can be separated. The aromatic-sulfur compound portion extracted from LCCO contains about 80 to aromatics, alkylbcn- Zenes and methyl and dimcthyl naphthalenes, 30 to 50% dicyclics and 1.8 to 2.5% sulfur, and boils over the range of 400 to 500 F. This aromatic-sulfur compound mixture can be deme-thylated by the process of this invention to produce a mixture of hydrocarbons of which 11- 35 to 50 volumepercent is 95 to 100+ octane gasoline and 7 to 15 volume percent is naphthalene.
The reformer distillation bottoms generally obtained as hereinbefore described contains about '70 to 75% methyl naphthalenes including mono-,diand tri-methylnaphthalenes. These bottoms can be used in the process of this invention or they can be distilled taking about 80 volume percent as an overhead fraction. This fraction will contain 85 to 90% of the mixture of methylnaphthalenes originally in the bottoms. head fraction boi-ls over the range of 448 to 550 F. and has the following ASTM distillation characteristics:
ASTM distillation: Degrees Fahrenheit Initial boiling point 448-458 10% 464-472 470-476 474-478 40% 476-480 480-484 1.. 483-488 488-490 494-496 506-508 Final boiling point 544-550 One typical overhead fraction h-as a gravity of 10.2
Such an overinstead of nameare the initial boiling points of the degrees API and contains on a volume basis: Alkylbenzones, 6%; naphthalene benzenes (tetralins, indanes, etc.), 7%; dinaphthalene benzenes, 2.0%; naphthalene, 1%; C n aphthalenes, 46% C naphthalenes, 26%; and C naphthalenes, 12%, the latter three being monoe, diand tri-methylnaphthalenes. Y
Mixtures of methylnaphthalenes such as mixtures of mono-methyl-naphthalene isomers, mixtures of 'dimethy-lnaphthalene isomers, mixtures of monoand di-methylnaphthalene isomers, etc., ho-wever obtained can be employed as methyl aromatic feeds for the process of this invention to produce naphthalene.
An embodiment of the process of this invention as illustrated in FIGURE 2wil-l be described in detail with respect to the preparation of naphthalene of 98+ percent purity (779.4 C. freezing point) from a reformer bottoms containing 85 volume percent (mono-, diand rtri-) methylnaphthalenes.
Referring to FIGURE 2, feed for thermal hydrodemethylation prepared by starting with 1351 gallons per hour reformer bottoms at 100 F., taken from storage (not shown) source through conduit '101 by pump i102 and discharged through conduit 103 to the upper portion of sponge oil absorber 104 which may be a packed column or a column filled with'fr-actionati on trays. Sponge oil absorber is operated at 210 pounds per square inch absolute (p.s.i.a.) and about 131 F. at the bottom with vapors leaving at about 105 F. To the lower portion of sponge oil absorber 104 is fed by conduit 16 7 a mixture (6032 pounds per hour) of flash gas streams obtained as hereinafter described. In sponge oil absorber about 95% of the benzene in the mixture fractions in F.
TABLE I Component Stream 103 Stream 106 Stream 107 Benzene 7. 46 9. 73 0.99 1. 24 0. 42 0. 51 0. 12 0.15
The mixture of flash gases in conduit 167a in addition to the normally gaseous components named in the tail gas composition also contains benzene.
The liquid flowing in conduit 107, 1520 gallons per hour at 131'F., passes through heat exchanger 108 and conduit 109 to the central portion of feed prefractionator 110 operated at a bottom pressure of 9.4 p.s.i.a. and 550 F. Additional heat is supplied by reboiler 114 by pass- 7 ing a portion of the liquid at 530 F. in the lower por- .that is mainly materials normally boiling at 550 F. and
above (81 mole percent) with the remainder boiling normally from 505 to 550 F., are removed as bottom fraction from feed prefractionator 110 at 231 gallons per hour by pump 112 through conduit 111. The fraction "taken overhead passes through heat exchanger 108 and pentane and butane are recovered. The tail gas containing about 68 mole percent hydrogen is withdrawn from the top of sponge oil absorber 104 by conduit 105 and discharged preferably to a fuel gas header. The tail gas also contains about 20 mole percent methane, about 8 mole percent ethane and lesser amounts of other C to C paraflins and olefins. The rich oil bottom-s from sponge oil absorber are withdrawn at 131 F. through conduit 106 and combined with 35 gallons per [hour liquid bottoms from separator 192 flowing in conduit 198 forming the liquid stream flowing in conduit 107, The composition, in moles per hour (m./i1.), of the liquid streams in conduits 103, 106, and 107 is shown in Table I wherein Components include only materials boiling from benzene and higher and wherein those I is cooled from 428 F. to 396 F. by preheating the stream in conduit 107 by indirect heat exchange. The partially cooled overhead fraction in conduit 117 is further reduced in temperature by cooler 118 to 125 F. and thence flows through conduit 119 to reflux drum 120 operated at 2.8 p.s.i.a. About 805 gallons per hour of condensate are returned as reflux by pump 124 to feed 'prefraction-ator 1'10 and #1 289 gallons per hour are removed by pump 127 through conduit 128 as feed for hydrodemethylation. This feed contains about 4 mole percent parafiins and olefins, 13 mole percent benzene and 80 mole percent aroma-tic materials, normally boiling above 450 F. I
The mixture of hydrocarbons containing methylnaphthalenes attained in the foregoing manner and a hydrogen feed are preheated by reactor eflluent and combined before entering furnace reactor 140in the following manner. The hydrocarbon feed (4 mole percent paraffins and olefins, 13 mole percent benzene and 80 mole percent other aromatics) at 1289 gallons per hour is pressurized by pump 127 and discharged through conduit 128, heat exchanger 129 wherein heat is extracted from reactor efiluent in indirect heat exchange entering heat exchanger 129 at about 590 F. and leaving at about 400 F. heating the feed hydrocarbon from F. to 510 F., a transfer of about 2.14 10 B.t.u./hr. The partially heated-feed hydrocarbon leaves heat exchanger 129 by conduit 130. Hydrogen feed at 3,385 pounds per hour (about 583 total moles), containing about 84.5 mole percent hydrogen, 6.7 mole percent methane, 6.5 mole percent ethylene, 2.3 mole percent propanes, is
13 taken from storage source (not shown) at 145 p.s.i.a. and 100 F. by gas transfer line 132 by compressor 133 and pressurized to 920 p.s.i.a. which increases the temperature of the hydrogen food to 300 F. This pressurized hydrogen feed is discharged to gas transfer line 134 and flows in indirect heat exchange with reactor eflluent in heat exchanger 135. The reactor effluent is thereby cooled from about 750 F. to about 590 F. While the hydrogen feed is heated from about 300 F. to about 620 F. The further heated hydrogen feed leaves heat exchanger 135 by gas transfer line 136. The partially heated feed flowing in conduits 130 and 131 and heated hydrogen feed flowing in gas transfer line 136 are combined in gathering conduit 137 (temperature of the mixture is about 560 F.). The mole ratio of hydrogen to hydrocarbon feed is about 7.75 to 1 or 7260 standard cubic feet of hydrogen feed per barrel of hydrocarbon feed.
The mixture of hydrogen feed and hydrocarbon feed flowing in gathering conduit 137 enters heat exchanger 138 wherein by indirect heat exchange with the hottest reactor effluent the mixtures of hydrogen-hydrocarbon feeds (reactor charge stock) is heated from about 560 F. to about 815 to about 895 F. depending on the operation of furnace-reactor 140. When the operation of the reactor section of furnace-reactor is isothermal, the reactor inlet and efiuent temperatures are about 1250 B; when the operation is adiabatic the inlet temperature is 1160 F. and the outlet temperature is 1325. Hence the lower temperature (about 815 to 820 F.) of the preheated charge stock results from indirect heat exchange with 1250 F. reactor efliuent and the high temperature (about 895 F.) results from indirect heat exchange with 1325 F. reactor efiiuent. Y
The reactor charge stock at 815 to 820 F. (or 895 F.) flows out of heat exchanger 133 through charging conduit 139 and into furnace coil 144 in radiant sections 141 of furnace-reactor 140. Furnace coil 144, for example, may be constructed from 22 tubes 3.5 inch ID. by 4.5 inch D. to provide a pressure drop not exceeding about 30 psi. through the entire coil, preferably the pressure drop should be in the range of 20 to 22 p.s.i. The reactor charge stock is heated to 1250 F. (for isothermal reaction), or to 1160 F. for adiabatic reaction, in furnace coil 144 by fuel, for example, a portion of the absorber tail gas taken from conduit 105, charged to burners 143.
The reactor charge stock preheated to 1250 F. (or 1160 F.) flows to reactor coil 145 joined to furnace coil 144. For isothermal reactions at 1250 F., the flue gas from burners 144 is quenched with air added, for example, by air nozzles 146. For adiabatic reaction (inlet to reactor coil 145 is 1160 F. and outlet is 1325 F.) no air quenching of flue gas is required. If necessary, for close temperature control in the adiabatic type reaction, a portion of hydrocarbon feed (up to 15%) can is employed to extract heat from the reactor effluent to be used to supply a portion of the heat required by reboilers 114, 221 and 224, as hereinafter described, according to the unique method of this invention. The efl luent thereafter flows through heat exchanges 135 and 129 to preheat partially the hydrogen feed and hydrocarbon feed, respectively as hereinbefore described. The eflluent now cooled to about 400 to 435 F. is further cooled to 210 F. flowing through heat exchanger 156 which may be a box cooler.
The reactor efliuent, still a mixture of liquids and gases, at 210 F. and about 760 p.s.i.a. flows through conduits 157 and 158 to a high pressure separator system illustrated in FIGURE 2 by first high pressure separator 159 and second high pressure separator 163. There is also added to conduit 158 by way of conduit 178 benzene rich condensate from reflux drum 174 withdrawn therefrom by pump 177 and conduit 176. There flows into first high pressure separator 159 14,864 pounds per hour (143 gallons perhour of condensate at 120 F. is added to reactor efliuent) at 760 p.s.i.a. and 200 F. About 6235 pounds per hour of vapors are flashed from first high pressure separator 159 through vapor transfer line 160 to cooler 161 from which the cooled vapors, including some vapors of phthalene, flow through transfer line 162 to second high pressure separator 163. The benzene rich stream from conduit 178 is added to reactor effluent to keep naphthalene in solution in second high pressure separator operated at 750 p.s.i.a. and 110 F. The benzene rich recycle condensate itself contains naphthalene, most of which remains in the liquid in first high pressure separator 150. Advantageously condensation of most of the naphthalene and 200 F. in first high pressure separator 159 minimizes the quantity .of benzene-rich recycle and conserves heat. Materials vaporized at 110 F. and 750 p.s.i.a. from second high pressure separator 163 plus the uncondensed gases flow through vapor transfer line 167 to sponge oil absorber 104 Where 95% of the benzene is absorbed and recovered in the rich oil bottoms Withdrawn by conduits 106 and 107.
Liquids from each of high pressure separators 159 and 160 are withdrawn therefrom by conduits 164 and 165 respectively and collected in conduit 166. About 1076 and 1132 gallons per hour flow in conduits 164 and 165 respectively. The composite naphthalene containing liquid at 195 F. flows to prenaphthalene tower 170 at the ture of about 510 F. maintained by reboiler 221 probe taken from conduit 126 and added to a portion of reactor coil 145 near the inlet thereto as hereinbefore described.
Conveniently, the reactor coil 145, constructed as hereinbefore described, may be 1350 feet of 4.5 inch O.D. tubes in series to remove about 2 10 Btu. per hour heat of reaction to maintain isothermal conditions in the thermal hydrodemethylation reaction.
From reactor coil 145 the eflluent at 1250 F. (or 1325 F.) at 13,826 pounds per hour total of liquids and gases flows through effluent conduit 148 to heat exchanger 138, thence through conduit 149. When the efliuent is at 900 F. after heat exchanger 138, about 43-44% is bypassed heat exchanger 151 by conduit 150 so that when mixed with the remaining 57-56% passing through heat exchanger 151 the mixture is at 750 F. When the effluent is at 1015 F. (as from adiabatic operation) after heat exchanger 138, all of the 1015 F. efiiuent inconduit 149 passes through heat exchanger 151. Heat exchanger 151 vides an overhead vapor stream withdrawn from the top of prenaphthalene tower 170 by vapor transfer line 171 which conducts the vapors to cooler 172'from which the condensate is withdrawn to reflux drum 174 operated at 20 p.s.i.a. and F. by conduit 173. A portion of this condensate is removed for recycle to high pressure separator 159, the remaining portion is withdrawn by pump 180 through conduit 179 to provide a reflux rate of about 3 to 1 in prenaphthalene tower and the remainder is cooled to 100 F. by cooler 183 and sent as light aromatics to storage. The light aromatics stream contains about 55 mole percent benzene, about 9.4 mole percent toluene and about 8.5 mole percent naphthalene, together with about 5.9 mole percent paraflins and olefins, and the remainder being aromatics. The light aromatics product stream amounts to about 130 gallons per hour, and are excellent as gasoline blending component.
Vapors from reflux drum 174 pass through vapor transfer line to knock-out drum 185 from which liquids are recycled to reflux drum 174 by conduit 186.
to 55 p.s.i.a. and 204 F. cooled by cooler 190 to 110 F. and 50 p.s.i.a. and separated into gas and liquid in separator 192 from which the liquids are'withdrawn through conduit 198 to be added to bottoms from sponge oil absorber 104. The gases from separator 192 leave the top thereof by vapor line 113 and are taken by compressor 194, pressurized to 210 p.s.i.a. at 241 F., cooled by cooler 196 to 110 F. and added by line 197 to vapors in vapor transfer line 167 flowing to the lower portion of sponge oil absorber 104. i
Bottoms stream from prenaphthalene tower 170, enriched in naphthalene, about 87 to 88 mole percent naphthalene, is taken by pump 201 through conduit 200 at about 6600 gallons per hour andcharged via conduit 202 to naphthalene recovery tower 203 where naphthalene product of about 79.4 C. freezing point is taken overhead. Naphthalene recovery tower may be operated at p.s.i.a. with a bottoms temperature of 550 F. provided by reboiler 224 and a ratio of reflux to netoverhead of about 3 to 1. 643 gallons per hour of naphthalene product are removed to storage via product line 214. Naphthalene vapor condenser .205 is preferably a box cooler where the overhead vapors can be readily subcooled to 250 F. with water inlet of 115 F. and outlet of 180 F. V
Bottoms from naphthalene-recovery tower 203 are withdrawn by pump 208 through conduit 207 and sent to storage via conduit 209. 7 These tars can be combined with tars in conduit 116 and sent to storage and used, for example, for blending heavy fuel oils.
The naphthalene recovery tower may be operated at atmospheric pressure at various tower top temperatures depending upon thepurity of naphthalene product desired. For example, a 78.6 C. freezing point'product can be recovered at 375-440" F. in yields of 65.2% based on the feedstock of 80% overhead fraction of reformer bottoms, a 78.45 C. freezing point product can be recovered at 415 to 435 F. in 65.0% yield on the same basis or a 79.75 Q-freezing point product can be recovered at 422- 'to 426 F. in a 47% yield on the same basis. The temperature ranges in F. are the tower top temperatures.
-Reboiler heat supply requirements for feed prefractionator tower 110, prenaphthalene tower 170 and naphthalene tower'203 total about 12.8 to 12.9 10 B.t.u./hr. Of this heat requirement a minimum of about 1.6 10 B.t.u./hr. can be extracted from reactor effluent by heat exchanger 151 when isothermal conditions are maintained in reactor coil 145. The remaining heat is supplied by reboiler fluid turnace 250 wherein, for example, a portion of the tail gas from sponge oil absorber 104 can be burned as fuel charge to the burner 249. This heat supply system comprises reboiler fluid furnace 250, fluid surge drum 230,
circulatory pump 232 and transfer conduits. The fluid flowing in this system can be any fluid which does not thermally decompose at 650 to 700 F. Such fluids com- ,phenyls, anzentectic mixture of diphenyl and diphenyl oxides such as .Dowtherm heat transfer medium sold by Dow Chemical Company, and other heat stable heat transfer mediums. i i
In this reboiler heat service system, reboilers 114, 221, and 224 shown with their respective fractionating towers 110, .170 and 203, are the same as reboilers 114, 221 and 224 shown in the reboiler heat service system of FIGURE 2. A hydrocarbon oil mixture thermally stable at temperatures up to 700 F. and having an API gravity of 30U is circulated in the reboiler heat service system at a total rate of 348,000 pounds per hour by circulatory pump 232 drawing the oil from surge drum 230 through conduit 231 and discharging through conduit 233. The oil in surge drum 230 is at 590-595 F. and about p.s.i.a. A substantial portion of the oil, 88 to 90%, is'taken from conduit 233 by conduit 234 and passed through coil 236 in furnace 250. The 10 to 12% of the oil at 590595 F. is passed through conduit 233a to heat exchanger 151 in indirect heat exchange with reactor effluent after it has preheated the reactor charge stock. The oil is heated 'about 63 F. in heat exchanger 151 and then flows through conduit 235 to be mixed-With the oil passing through furnace coil 236 discharging via conduit 237. Themixed oil has a temperature ofsabout 655? F. This oil flows in hot oil supply line 238 to each of reboilers 114, 221 and 224where itis cooled to 600 F., 550 F. 'fand600 F. respectively, providing 120,000, 53,000 and 143,000 B.t.u./hr. respectively. Only about 92% of the circulating. oilis needed to pass through these reboilers, but valved conduit 252 is provided to regulate the flow of by-pass oil therethroughto adjust the hot oil flow as required by the heat load demand. After supplying heat to reboilers 114, 2 21 and 224, the oil is gathered in gathering conduit 243. and returned at about 590595 F. to surge drum 230. I r i The further illustrate the process or" this invention there is employed as the feed charge the 10.2 degree API overhead fraction hereinbefore described containing 46% C naphthalenes (monomethyl), 26% C naphthalenes (ethyl and dimethyl) and 12% C naphthalenes (mainly trimethyl). This methyl aromatics feed is employed at yarious temperatures, hydrogen to methyl aromatics ratios, (s.c.f./b.) pressures (p.s.i.g.) and space velocities '(V /hL/V The results of several runs are shown in Table II. In this Table I1 s.c.f./ b. is used to indicate standard cubic feet per barrel of methyl aromatics feed charge.
TABLE 11 Hydrodemethylation of aromatics concentrate of reformer bottoms Isothermal reactor N aphthalene product Coke yield, Run Hydrogen weight percent N 0. ratio SOF/B Pressure, Tempera- Space velocon feed charge Yield-weight Mole percent p.s.i.g. ture, "F ity, V /hrJV percent on feed on alkyluaphcharge thalenes 5, 800 1, 147 1. 0 57. 7 77. 8 9,800 800 1, 194 1. 0 0.010 57. 6 77.7 8, 600 800 1,192 1.0 0. 010 59.0 79. 5 5, 200 800 1, 199 1.0 0.010 60. 0 82.2 11,000 600 1,196 1.0 0.022 48. 4 65. 2 5, 700 600 1, 194 1.0 0.005 55. 7 75.0 6, 900 800 1, 205 1.0 Nil 63.0 85.0 7, 100 800 1, 239 3. 0 40. 9 54. 7 7, 000 800 1, 240 6.0 33. 9 45. 1 6, 900 800 1, 194 1. 0 0. 073 57. 7 77.8 7, 300 800 1,196 1. 0 0. 015 64. 6 87. 3 6, 700 800 1,120 0.5 59. 5 80.3 6, 400 800 1, 082 0. 5 57. 1 77. l 12, 200 800 1, 181 0. 5 0. 041 65.7 89.3
1 None detected.
2 In this run the feed charge is 75 weight percent aromatics concentrate and 25 weight percent postnaphthalene traction and yields are based on total mixture.
The separation system for recovering light aromatics and naphthalene can be carried out at different temperatures and pressures, if desired, as will be appreciated by the skilled design engineer.
What is claimed is:
1. In a process for the thermal hydrodemethylation of methyl substituted aromatics in the substantial absence of demethylation catalyst by subjecting to elevated temperatures in a demethylation zone a charge stock comprising hydrogen-eontaining gas and methyl aromatics feedstock at hydrogen pressures in the range of 450 to 1000 p.s.i.g., the improvements for the thermal hydrodemethylation characterized by the steps comprising: heating the charge stock by indirect heat exchange with effluent from said demethylation zone in the sequence wherein demethylation zone efliuent firs-t heats said charge stockto a temperature within 300 to 400 F. of the demethylation temperature and the partially cooled demethylation effluent then heats separately the hydrogencontaining gas and said methyl-aromatics feedstock which are then combined to form the charge stock; introducing the partially preheated charge stock at a volume hourly space velocity of feedstock of at least 0.5 to a combination of a separate heating zone and a separate demethylation zone wherein the charge stock is heated to a temperature in therange of 1100 to 1340 F. in the radiant zone of said heating zone and is maintained at a temperature in the range of 1100 to 1340" F. in said demethylation zone; and withdrawing demethylation zone effluent for said preheating sequence.
2. The process of claim 1 wherein the charge stock flows in the heating zone in parallel free flow paths and in the reaction zone in series flow path.
3. The process of claim 1 wherein the partially preheated charge stock is heated in the heating zone in parallel free flow paths to a temperature in the range of 1200 F. to 1340 F. and is maintained under isothermal conditions at a temperature in the range of 1200 F. to 1340 F. in the reaction zone in series flow path.
4. The process of claim 3 wherein the heating zone is in the radiant zone of fuel burned to supply heat thereto and the flue gas therefrom is quenched with air to remove heat of reaction from the reaction zone to provide iso thermal operation of the hydrodemethylation reaction.
5. In a process for the thermal hydrodemethylation of methyl substituted aromatics in the substantial absence of demethylation catalyst by subjecting to elevated temperatures in a demethylation zone a charge stock comprising hydrogen-containing gas and methyl aromatics feedstock at hydrogen pressures in the range of 450 to 1000 p.s.i.g., the improvements for the thermal hydrodemethylation characterized by the steps comprising: heating the charge stock by indirect heat exchange with eflluent from said demethylation zone in the sequence wherein demethylation zone effluent first heats said charge stock to a temperature within 300 to 400 F. of the demethylation temperature and the partially cooled demethylation efiluent then heats separately the hydrogen-containing gas and said methyl aromatics feedstock which are then combined to form the charge stock; introducing the partially preheated charge stock at a rate to provide a volume hourly space velocity of feedstock of at least 0.5 to a combination of a separate heating zone and a separate demethylation zone wherein the charge stock is heated to a temperature in the range of 1100 to 1340 F. by flowing in parallel flow paths in the radiant Zone of fuel burned in said heating zone and is maintained under isothermal hydrodemethylation conditions at a temperature in the range of 1100 to 1340" F. by flowing in series flow path in said demethylation zone by indirect contact with flue gas from said heating zone quenched with air; withdrawing demethylation zone eflluent from said preheating sequence, and flashing from the resulting partially cooled demethylation zone effluent a gaseous mixture containing hydrogen and methane split oii: during hydrodemethylation reaction, and
burning a portion of said gaseous mixture as fuel for said heating zone.
6. The process of claim 5 wherein the portion of demethylation zone effluent after removal of said gaseous mixture containing hydrogen and methane is fractionated to recover demethylated aromatic product.
7.,The process of claim 6 wherein heat is extracted from demethylation zone effluent prior to preheating the hydrogen-containing gas and the feedstock to provide a portion of the heat for the fractionation of the portion of the demethylation zone efiluent after the gaseous mixture containing hydrogen and methane is removed.
8. An integrated system for thermal hydrodemethylation of methyl aromatics in the presence of hydrogen comprising a furnace-reactor containing a heating zone and a reaction zone and a means for supplying heat to said heating zone to provide therein a radiant zone, and interconnected means for passing hot gases from said heating zone to said reaction zone, parallel flow paths in said radiant zone connected to demethylation zone comprising a series flow path in said reaction zone, charging conduit for supplying charge stock comprising hydrogen and methyl aromatics to said parallel flow paths, a charge stock preheater, a conduit for transfer of eflluent from said series flow path to said charge stock preheater for indirect heat exchange therein, a preheater for heating pressurized hydrogen-containing gas, a methyl aromatics feedstock preheater, conduit for transfer of said effiuent partially cooled from said charge stock preheater sequentially to each of said preheaters for heating pressurized hydrogen-containing gas and methyl aromatics feedstock, separate conduits for separately introducing pressurized hydrogen-containing gas and feedstock for indirect heat exchange to their respective preheaters, gatherin conduit for combining said preheated pressurized hydrogencontaining gas and feedstock to provide said charge stock for indirect heat exchange in said charge stock preheater, means for flashing gas from the cooled etfluent, conduit for transfer of cooled efiluent from said preheaters for pressurized hydrogen-containing gas and feedstock to said flash chamber, a gas cooler, a gas-liquid separator, cor1- duit for conducting gases from the top of said flash chamber through said cooler to said gas-liquid separator, conduit for removing gas from said gas-liquid separator, a fractionation system for recovery of demethylated aromatics, and conduit for charging degassed effluent to said fractionation system from the bottom of each of said flash chamber and said gas-liquid separator.
9. The integrated system of claim 8 which includes an intermediate heat exchanger for removing :by indirect heat exchange heat from partially cooled eifluent from said series flow path intermediate said charge stock preheater and preheaters for pressurized hydrogen-containing gas and feedstock, conduit for transfer of said partially cooled efiluent from said charge stock preheater to said preheaters for pressurized hydrogen-containing gas and feedstock, and conduit from at least one reboiler in said fractionation system to said intermediate heat exchanger to supply a portion of reboiler heat for said fractionation system.
10. An integrated system for thermal hydrodemethylation of toluene to produce benzene comprising a furnacereactor containing a heating zone and a reaction zone and means for supplying heat to said heating zone to provide a radiant zone therein, parallel flow paths in said radiant zone interconnected to demethylation zone comprising a series flow path in said reaction zone, charging conduit for supplying charge stock comprising hydrogen and toluene to said parallel flow paths, a charge stock preheater for indirect heat exchange therein, a conduit for transfer of effluent from said series flow path to said charge stock preheater, a preheater for heating pressurized hydrogen-containing gas, a toluene preheater, conduit for transfer of partially cooled effluent from said charge stock preheater first to the hydrogen-containing gas preheater and then to the toluene preheater, separate conduits for introducing pressurized hydrogen-containing gas and toluene for indirect heat exchange to their respective preheaters, gathering conduit for combining preheated pressurized hydrogen-containing gas and preheated toluene to provide said charge stock for said charge stock preheater and said parallel flow paths, a cooler, a flash chamber, conduit for transfer of cooled effluentfrom said toluene preheater through said cooler tosaid flash chamber, a firstgas cooler, a liquid gas separator, conduits-for transfer of gas from the top of said flash chamber through said gas cooler to said liquid-gas separator, con-duit for Withdrawing gas from the top of said liquid gas separator and for charging a porti-onvof the removed gas as fuel to the heating zone of the furnace-reactor, a fractionating system including a stripping tower, a benzene fractionator and a toluene fractionator, conduit for the transfer of liquid from the bottoms of the liquid-gas separator and the flash chamber to said stripping tower, a second gas cooler, a second liquid-gas separator, conduit for transferring gases from the top of said stripping tower through said second gas cooler to said second liquid-gas separator, conduit and compressor for transferring and pressunizing gas from the top of said second liquid gas-separator to said first, gas cooler, conduit and pump for recycling als reflux liquid from the bottom of said second liquid-gas separator to the upper portion of said stripping tower, pump and conduit for transfer of liquid from bottom of stripping tower to feed to benzene fractionator, reflux system'forsaidbenzene fractionator with pump and conduit for withdrawing benzene product from benzene reflux, pump and conduit for transfer of bottoms from benzene fractiona tor as feed,
demethylated aromaticswherein a demethylation stock 7 containing hydrogen and methyl aromatics feedstock and the charge stock components are preheated toua temperature below demethylation temperature and :wherein de:
pressurized hydrogen-containing'gas and feedstock, a storage vessel for fractionation reboiler heat transfer liquid, a means for heating a major portion of said heat transfer liquid; conduit for sequentially passing said effluent containing demet-hylated aromatics through said charge stock preheater and thereafter through said heat extractor and said preheaters for pressurized hydrogen-containing.gas and feedstock, conduit and pump for transferring heat ex change liquid from said storage vessel, conduit for charging a major portion of said heat exchange liquid to said means for heating the major portion of the heat exchange liquid, conduit for transfer of the remaining minor portion of: said heat transfer liquid through said heat extractor, conduit for withdrawing saidlrnajor portion of the heat exchange'liquid from said heating means therefor, connections for recombining said major and said minor portions of said heat exchange liquid in a gathering conduit,
. and lines for conducting the recombined heat exchange liquid through the fractionation reboilers to said heat exchange liquid'storage vessel. 7 l
12. In an integrated process for producing naphthalene and light aromatics for gasoline blending from methyl 1 n-aphthalenes enriched fraction of said reformer bottoms,
fractionating the effluent mixture from hydrodemethylation to remove a hydrogen-methane gaseous fraction, said lightaromaticsfraction and product naphthalene; the improvement comprising stripping said catalytic reformer bottoms with the hydrogen-methane gaseous fraction to obtain. .a methyl naphthalene enriched fraction of said catalytic reformer bottoms, recycling a portion of benzene from the hydrodemethylation effluent to the fractionation methylated aromatics product :is recovered from saidefliuent in a fractionation system containing fractiona-tor reboil= ers heated with a heat'transfer liquid, an integratedheat exchange system for preheating charge stock and thecomponents thereof and for supplying a portion of the heat for said fractionation system which comprises a charge stock preheater, a heat extractor, preheaters for heating of hydrogen-methane gas from said effluent thereby keeping naphthalene in solution and recycling a portion of benzene to said methyl naphthalenes enrichment.
References Cited by the Examiner UNITED STATES PATENTS 1,872,011 8/ 32 Russell 208--l07 1,934,056 11/33' Gomory 208107 2,907,800 10/59 Mertes 260-672 ALPHONSO D; SULLIVAN, Primary Examiner.
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|US8070855 *||Dec 19, 2008||Dec 6, 2011||Strickland Michael L||Methods and apparatuses for reducing emissions of volatile organic compounds from pumps and storage tanks for VOC-containing fluids|
|US8679230||Oct 27, 2011||Mar 25, 2014||Michael L. Strickland||Reducing emissions of VOCs from low-pressure storage tanks|
|DE1301408B *||Dec 4, 1965||Aug 21, 1969||Universal Oil Prod Co||Verfahren zur Umwandlung eines alkylaromatischen Kohlenwasserstoffoeles mittels Entalkylierung|
|U.S. Classification||585/402, 585/483, 585/484, 208/107|
|International Classification||C07C15/00, C10G35/00, C10G47/00|
|Cooperative Classification||C07C15/00, C10G47/00, C10G35/00|
|European Classification||C10G35/00, C07C15/00, C10G47/00|