US2320147A - Aromatization - Google Patents

Description

TIMER@ C104? (F) May 25, 1943. E. T. LAYNG Erm. ARoMATIzATIoN Filed sept. 1:5, 1959 Sheets-Sheet 1 May z5, 1943;

E. 1'. LAYNG ErAL ARQMAT1ZATION Filed sept.v 1s. 1939V 7 Sheets-Sheet 4 Moz 13,1710 v Y E. T. LAYNG EI'LA May 25, 1943.

AROMATIZATION File@ Sept. 13. 1939 May 25', 1943. E. 1'. I YNG ErAl.

AR'oMATIzATIoN Filed sept. 1:5, 1939 '1 Smets-sheet' e E. T. LAYNG` ETAL AROMATIZATION Filed Sept.' 13, 1939 May 25,.- 1943.

Patented May 25, 1943 AROMATIZATION Edwin T. Layne and Louis c. Rubin, Jersey city, v N. J., assignorsto The M. W. Kellogg Co., Jersey City, N. J., a cpi-poration of Delaware Application `September' 13, 1939, Serial No. 294,784

(Cl. 26d-668) 16 Claims.

This invention relates to the.l conversion, o'f hydrocarbon stocks boiling within, and in some cases within and somewhat above, thegasoline l boiling point range to high grade motor fuel by catalytic aromatization under critically defined conditions.

It is known that by the use of certain catalysts naphthenic hydrocarbons can be converted into aromatic hydrocarbons with or without the presence of hydrogen. It is also known that at approximately atmospheric pressure inthe absence of hydrogen, paraflinic hydrocarbons can be converted into aromatic hydrocarbons by the use of certain catalysts. It has also been known for years that destructive hydrogenation under Thigh pressures and inv lthe presence of large amountsvof added hydrogen will produce motor fuels which contain a limited amount of arolmatics and which have a relatively high octane number from stocks boiling well above the gasoline boiling point range, but such processes have not made appreciable commercial headway because of the expense of building and operating a trol a largenumber of operating variables with -inl closely and critically defined limits; for income apparent as the description thereofproceeds.

We have found that in order to overcome the disadvantages of the prior art and achieve the objects above enumerated, it is necessary to constance, we nd that the severity factor, which term we will later denne, must be great in order to achieve the production oi aromatics from parailinic hydrocarbons and such oleflnic hydrocarbons, if any, as may be present. At the same timelwe ndvthat the severity factor must not be excessive or adverse inuences come into play. Concurrently with control of the severity factor, we also nd that pressure in conjunction with added hydrogen lriustbe used to decrease coke formation and increase octane number 'of the hydrogenation unit which operates at extremely,`

high pressures (not to mention the expense of producing or supplying the necessary hydrogen) n and because of the undesirable increase in vola, tility and the production of hydrocarbon gases.

An object of our invention is to obtain theI advantages of all of these previous processes without suffering their inherent disadvantages by means. of an improved endothermic dehydroaromatization process.

Prior to our work no commercially practicable process for the conversion of aliphatic, particularly paraillnic, hydrocarbons boiling within and slightly above the-motor fuel range to motor fuels of high aromatic content has been developed. In

the course of a very great amount of experimental work we have found that under critically defined conditions stocks boiling within, and sometimes slightly above, the gasoline boiling point range can be converted into unusually high yields of aromatic hydrocarbons boiling within.

the gasoline range, at the same time accomplish-y ing a number of advantages not .achieved bythe p'rior art.

It is an object of our invention to provide a.' process for the conversion of aliphatic hydrocarbons to aromatic hydrocarbons. It is a further object of our invention to provide a process for the conversion of a stock boiling approximately within the gasoline boiling point-range, and rich extremely small amounts of coke and gives long4 catalyst life. Other and more detailed objects, advantages and uses of our invention will besuch as we have devised. Further we find that the pressure must be critically controlled since if this is not done light parafllns are produced which, if gaseous, reduce the liquid yield, and if liquid and having less than six carbon atoms cannot be aromatized thus limiting the octane number to the relatively low value characteristic of paramns.

v In spite of the factv that our process is v ery definitely one of dehydrogenation and aromatization ratherfthan hydrogenation, we 'find it `very important, likewise in conjunction'with the other noted factors, to use critically controlled amounts loi hydrogen which are in general considerably less than are used in hydrogenation processes. The use of hydrogen cuts down coke formation and greatly increases the length of run possible.

u Further, the use of hydrogen up to a certain low fig-ure gives an important improvement in the octane number of the product but beyond that 'figure the octane number begins to drop off. Thus Y .A we have found that there is an optimum' amount I fer the oxides of the metals of the left-hand column of group VI of the periodic table, particu- 'larly chromium, molybdenum and tungsten but we can also use other metallic oxides and other metallic compounds, particularly oxides of the .metals of the left-hand columns of groups IV and V of the periodic table such as titanium. cerium, thorium and vanadium. Moreover. while these catalytic oxides can be used alone or on various y supports including magnesla, we find it very highly preferable to utilize them on alumina, particularly an activated alumina r on alumina gel, as a support and in general the catalytic oxide or other catalytic compound should be the minor constituent, usually from 1 tc 40% by weight of the total catalyst including the support, although the optimum percentage varies, cf course, with the catalyst used. It will also be apparent that mixed catalysts can. be used, for instance a mixture of chromium oxide and molybdenum oxide alone or preferably on an alumin'a support, and in this case the active catalytic oxides should be from about 1 to about by weight of the totalvcatalyst. Another catalyst which can be fused .is magnesium. chro-mite either alone or on a suitable support, preferably alumina. In fact, any aromatization catalyst or material that promotes the dehydrogenation and cyclization of aliphatic hydrocarbons can be used. As those skilled in the art know, the emciency of a given catalyst is largely determined by its method of preparation and con sequently some catalysts give better results than 'others under a. specic set of conditions. Such .differences obviously can be overcome to some extent by suitable slight alterations in the operating conditions within the limits taught inthe present specification. However, of all these catalysts we greatly prefer to use a catalyst com? prising molybdenum oxide supported on alumina, particularly activated alumina, and the molybdenum oxide should most advantageously constitute from about 2 to about 10% by weight of the total catalyst since this has been found rto include a very sharp optimum range which gives the best results. Larger amounts of molybdenum oxide can, however, be used.

We greatly prefer that our catalyst be substantially sulfur free since sulilde catalysts as contrasted with oxide catalysts tend to degrade the stock to less valuable light parans and lead to predominant hydrogenation rather than predominant dehydrogenation, i, e. hydrogen consumption instead of hydrogen production. These' sulfide catalysts also result in increased coke production.

Our preferred catalyst used in the experiments to be described in connection with Tables l.l

through 5 which will be given later was a molybdenum oxide on activated alumina catalyst and more particularly was made in accordance with the following procedure:

Dissolve 165 kilograms of ammonium para molybdate in sufficient distilled water tof yield 150G liters of solution. Place 1500 kilograms of granular activated alumina in an apparatus which can be evacuated and add the molybdate solution. Agitate the mixture and then apply a vacuum pump and reduce the pressure to 30-46- millimeters of mercury. Then allow the pressure to rise to atmospheric. Lower the pressure a second time to the same level and then allow it to rise again to atmospheric pressure. Repeat this procedure a third time and then drain the remaining liquid from the impregnated alumina. Air-dry the latter on screens or other suitable containers using layers of about one inch in depth. At the end of this time place the dried material in a furnace in a suitable container and heat it at a temperature of l20 F. for one hour. Cool to atmospheric temperature and store in closed containers until ready for use.

While it can be calculated from the quantity of solution absorbed that the catalyst should contain 3.75% M003 by weight, actual analysis shows -the presence of about 6% of molybdenum oxide,

this difference resulting, no doubt, from preferential adsorption of the molybdenum compound from the solution. In any event the dierence in effect between a catalyst containing 3.75% of M003 and one containing 6% is negligible.

As previously indicated, the oxide catalysts are preferred but in some cases, such as our preferred -molybdenum oxide catalyst, the oxide tends to be reduced in the process and in general instead of using the oxides of the metals of the left-hand columns of group IV, V and VI, the metals themselves can be used, preferably supported on alumina.

Our invention will be described further with reference to the accompanyingr drawings which form a part of this specification and in which:

Figure 1 is a plot of the area of severity factor, temperature and time factor in which we und it necessary to operate our process, the crosshatched area being preferred;

. Figure 2 is a plot of the relationship between v severity factor and octane number, which like Figures 3 to 7 is particularly applicable to a. particular charging stock and catalyst; but is representative of more general results;

Figure 3 is a plot of the relationship between the content ofv aromatic hydrocarbons in the j product from our process and octane number of l this product;

Figure 4 is a plot of the effect of increasing` severity factor on the amounts of aromatic and olefinic hydrocarbons in the product:

Figure 5 is a 4plot of the effect of increasing severity factor on the A. P. I. gravity of the produc z Figure 6 is a plot of the effect of the operating pressure on the amount of coke formed and on the octane number of the product;

Figure 7 is a plot of the effect of the mol ratio of hydrogen to charge on the amount of coke foizmed and on the octane number of the produc Figure 8 is aplot of the relationship between the operating variables severity factor and mol ratio of added hydrogen. and shows the area in which we nd it essential to work, the preferred area being cross-hatched;

Figure 9 is a. simplified flow diagram showing a stationary bed type of apparatus which can be used in accordance with our invention;

VFigure 10 is a simplified ow diagram of a moving bed type of apparatus which can be used in accordance with ourinvention;

Figure 11 is a detail of the apparatus for introducing catalyst and feed into the reaction chamber of Figure 10; and

Figure l2 is a Adetail of the apparatus for removing catalyst and products from the reaction chamber of Figure 10.

One of the important operating variables in our process is time factor which may be defined as the amount of time in hours required to put through the catalyst a volume of feed (measured as liquid) equal to the volume of the catalyst chamber, the volume of the catalyst chamber being the over-all volume of that portion of the chamber which is filled with the catalyst. In other words, time factor is the volume of catalyst space divided by the feed rate in volumes per hour. In determining this ratio, the term "volume of catalyst space obviously refers to .a space completely lled with the catalyst particles and -having a density corresponding to that which .reduced by extending the distance and increasing the voids between adjacent catalyst particles by mixing the active catalyst material,

with inert particles or otherwise, a suitable cor rection should be applied to compensate for the decreased concentration of active catalyst material in unit volume of catalyst space thus effected as will be obvious to those skilled in the art. We have found that the time factor should be between 0.1 and 25 and preferably between 0.2 and 20 hours.I Time factor or its reciprocal, the space velocity, is the important factor rather than contact time. Contact time, in the usual sense, is a calculated value intended to be the average length o'f time during which any given part of the charge is in contact with the catalyst. This is entirely unsatisfactory to calculate since it depends on factors which can only be estimated,

one of the most serious difficulties in4 figuring contact time being the fact that the time whichI is really signicant is the time an average molecule is present in an adsorbed lm on lthe catalyst. Calculated contact times have been found to be entirely insignificant and cannot be correlated with the results obtained in our process. This is particularly true because we chooseto y circulate hydrogen with the hydrocarbon charge.

Not only is time factor important but temperature is likewise limportant and should be at least 875` F. and preferably at least 890 F. On

the other hand, the maximum temperature is likewise significant and should be not over 1075" f F., If the temperature is lower than the minimum specified the reaction i's unsatisfactorily slow and does not produce the desired high yield of aromatic hydrocarbons from parainic hydrocarbons. On the other hand, if the maximum is exceeded an excessive amount of gas is produced. In other words, operating below the specified Alower limit tends to give a condition favorable to hydrogenation which is definitely not desirable vin our process. At the same time when the tem- ;erature is higher than the upper limit, thermal reactions become excessive.

When we speak of temperature in this specification and in the appended claims we refer to the temperature as measured in the catalyst bed unless some other meaning is indicated." The catalyst bed temperature to which reference 'is made is the weighted average temperature calculated from temperature measurements made at a plurality of representative points in the catalyst bed.

In addition to the importance of the time factor and temperature, the relationship between these two variables must likewise be maintained within critical limits. We refer to this relationship as severity factor and define the severity, S, as

S=T+1oo logm F Where T is the temperature in degrees fahrenheit and F is the time factor in hours. This equation defines severity in such terms that with a given charging stock and other process conditions being held constant, all runs using a given severity factor will giveapproximately the same octane number product. Since the high octane number of our product is due largely to its high the octane numbers mentioned herein, was 42.3.

" Table 1 Example number 1 Run number 912 Severity factor (S) 1067 Temperature (T) "F.1 947 Time factor (F) hrs-- 16' Pressure ..-lbs/sq. in..- 100 Mol ratio, Ha/charge (M) 8 I Catalyst holding time ....hrs 10 Product characteristics 2 Octane number 3 91.3 Yield per cent 3-.. 72.1 Gravity A.P.I 37.7 Aromatics per cent ..l 73.4 Olefins i-; do....- 7.6 Coke do 4..--'. 1.8 Difference between 50% distillation points of charging stock and product 2 F 28 0.90

content of aromatic compounds, it may be said that the content of aromatics is a function of severity factor and a certain minimum severity factor must be preserved in order to secure the beneficial results flowing from our process. With extremely active catalysts this minimum severity factor can be about 925 but in order to obtain the best results the minimum severity factor should be 950 or 975. At the same time if either the time factor or the temperature is too high,

unsatisfactory results are obtained and there is thus a critical maximum as well-v as a critical minimum on severity and this maximum is 1100 or, for best results, 1075. The area of temperature, time factor and severity in which we :Gnd it desirable and necessary to work in order to achieve our results, is shown inf Figure 1. The diagonal lines extending across this area are lines of constant severity and the numbers on them are the severities as above defined. The crosshatched area is that which is preferred.

One striking illustration of the results of high l' severity is shown in Table 1 iny which with a. severity of 1067 and the other conditions shown an octane number of 91.3 was obtained charging a stock with an octane number of 42.3. This is an octane number increase of 49 points compared with an increase of perhaps 10 or 12 points for an ordinary dehydrogenation process. The data shown in Table 1, like all the rest .of the data. in this specication, with one exception, were obtained using apparatus of the type shown in Figures 10, 11 and 12. As in all of Examples Nos. 1 through 12,\ the catalyst was the specific molybdenum oxide on alumina catalyst above described, and the charging stock was an East Texas heavy naphtha having an A. P. I. gravity of 50.3, containing approximately' 14% aromatics, 33% naphthenes and no olens, the remainder being largely parain hydrocarbons. This naphtha had an initial boiling point of 248 F., a 50% boiling point of 310 F. and an end point of 396 F. The original octane number measured by the Cooperative Fuel Research Motor Method, as were all Hydrogen produced, mols I-Iz/mol charge..

' carbons and no lighter hydrocarbons.

As percentage of charging stock.

is, typical of a large amount of data. Representative data illustrating this point are given in the following table. These data and other data given hereinafter were obtained with' the charging stock and catalyst previously described and the foot- 5 notes in connection with Table 1 are equally applicable to subsequent tables.

Tabe 2 Example number Run number. 912 887 902 909 927 l083A ]5 Severity factor (S) 1067 1037 1031 1006 985 945 'lemperature(T)F 947 940 971 946 894 947 Time factor (F) hrs. 16 8.08 4 4 8 0.958 Pressure lbs./sq. in... 100 100 100 100 200 `300 Mol ratio, 13g/charge 8 4 8 4 4 2. 8 Catalyst holding n ume hrs 1o 1o 1o 1o 10 212 .-0 Product characteristics:

Octane number. 9i. 3 87. 3 87.9 8l. 3 77. 6 73. fi Yield ercent 72. l 75.3 78. 2 82.3 83.9 87. l Gravity A. P. I.. 37. 7 41. 3 39. 7 42. 6 46. 6 48. 8 Aromatics percent 73. 4 65. 6 62. 9 48.1 40. 0 9,-) 0leins do. 7. 6 6.4 10. 1 10.9 10.0 Coke... d0 1.8 2.4 1.0 1.8 0.5 Difference between 50% distillation points oi charging stock and product, 2e 30 4o i5 Hydrogen produced, 31W

mois Hg/mol charge. 0.90 l. 0.91 0. 8A 1.8

l Fixed bed run. l i Length oi' run.

It will be seen from the foregoing table that octane number of product decreases rather rapidly with decreasing severity factor although yield increases somewhat at the same time. The exact quantitative relationship varies to some extent with the charging stock, the catalyst and with other operating conditions.

As previously mentioned, the high octane number of the product is due largely to the very high content of aromatic hydrocarbons and this is ils lustrated in Figure 3 which shows the content oi aromatic hydrocarbons as a function of octane number. This curve was plotted from a large number of runs using the East Texas heavy naphtha charging stock to which reference has previously been made. We can, of course, plot aromatic content against severity factor instead of against octane number and this is done in Figure i which was prepared largely from the data of Table 2. It is interesting to note that not only does our high severity result in very high aromatic content but it also results in low olefin content in the product as is also shown in. Figure 4. While with severity factors lower than those which we use the olen content of the productis often higher than that of the charge (assuming a charge low in olens), it will be seen that with the severity factors which we use the oleiin content is actually decreasing with increasing severity (or increasing octane number). The amount oioleiins relative to parans in the product for this particular naphtha tends to reach an equilibrium with the olens at 29% of the total olens plus parains so that as aro matics are produced in large quantities, thereby decreasing paraiins, olelns are decreased. Thus lwe have a dehydrogenation process which produces very little olefins and in fact converts oleilns, if present, as well as parailns to very large yields of aromatics.

The A. P. I. gravity of the product is a good rough index. of its content of aromatic hydrocarbons since aromatics have low A. P. I. gravities or, in other words, high specic gravities. Figure 5 shows an average plot of octane number versus A. P. I. gravity for products from a large number of runs including all those given in the examples in this specication and shows that with increasing octane number the gravity tends to drop of! and in any event the gravity of this stabilized product is less than, and for the more desirable severity factors usually at least 3 A. P. I. less than, the gravity of that part of the original charge having the same end point as the end point oi the product. As a, matter of fact, in spite ofv there being some material formed that boils below the initial boiling point of the charging stock, the gravity of the whole product in degrees A. P. I. is practically invariably less than that of the Whole charging stock (evenif the product contains material boiling above the gasoline boiling point range) when working under the conditions which We have determined. This is true even though increasing A. P. I. gravity ordinarily accompanies' increasing volatility thus partially offsetting the gravity effect resulting from the production of aromatics.

The low A. P. I. gravity of the gasoline made in accordance with our invention means that the customer gets more 'pounds per gallon as well as a better product.

While temperature, time factor and more particularly severity are very important in order to achieve the results of our process, still other factors are of very great importance.

For one thing pressure likewise must be kept within a critical range. have been described in the literature and in prior patents for carrying out aromatization reactions with some of the same catalysts which we have described but under atmospheric pressure and while the results are interesting as laboratory runs they are not of commercial interest due to the fact that the catalyst is very rapidly fouled by the deposition of coke and other carbonaceous matter so that the length of run in stationary bed operations or the catalyst holding time in moving bed operations is very uneconomically low under these conditions. We have found that notonly does the use of moderate pressures very markedly decrease the amount of coke deposition and'increase the length of run or, what amounts to the same thing, the catalyst holding time but the use f pressure, within certain limits and in the presence of hydrogen, increases the octane number of the product. The effect of increasing the pressure from atmospheric to 30l`pounds per square inch gage in a pair of comparable runs was to increase the octane number of the product lfrom 57.8 (atmospheric pressure) to 63.2 (30 pounds per square inch). While these octane numbers are low due to the fact that the tcm-A fer to use pressures of 50 to 300 pounds per square inch gage although good results can be obtained at from 30 to 375 or in some instances even 450 pounds per square inch gage. Naturally, the exact pressure Will vary depending upon the A number of processes increase inoctane number), each pair of data being under otherwise comparable conditions.

Figure 6 illustrates for typical conditions the eii'ect of pressure on the amount of coke and also on`octane number. These curves are plotted in part lfrom the data of the examples given in this specification and in part from additional data and trends. raising the pressure increases the octane numben at first, this increase attens off at about 200 or'250 pounds per square inch gage and finally tends to decrease in spite of the fact that increased pressure 'at a given time factor inevitably means an increase in the time during which a given particle of the charging stock is in contact with the catalyst.

While we prefer to charge straight run naphthas to our process, cracked naphthas can also be charged and in 'this event we pref r pressures of about 30 to 100 pounds per sql?? 'in `any event not over about 25 pounds per square inch since at a pressure of the order of `300 pounds per square inch there is aconsiderable tendency (with cracked stocks) to consume hydrogen instead of producingit.

Another 4very important considerationy in achieving the results of our process is careful control of the amount of hydrogen. We 4measure the amount of hydrogen in terms -of mol ratio vwhich appears to be the most significant measure. The mol ratio of hydrogenl is figured as down the formation of coke very markedlyvand thereby increases the length of runs or onstream perigdsto the point where regeneration is ot necessary at the frequent intervals which would otherwise be essentiaLThus lengths of` runs from 25 to 100 hours with the stationary 75 charge.

It will be noted that while e inch and catalyst bed can be used in our process or when using moving bed the catalyst holding time, i. e. the time the catalyst is in the reaction chamber, is of the same general magnitude. As applied to stationary bed work the term catalyst holding time can be used to refer to the length of run. "Followingthe period during which the catalyst is in the chamber (which may be very short or may be as long as about 100 hours) the catalyst must be regenerated which can be done by known means using air diluted with flue gas or using somel other oxygen containing gas as will later be described in connection with the flow diagrams.

While we achieve long catalyst .on-stream periods with satisfactory conversion as compared with prior art processes for aromatlzlng Table 3.

Example number Run number 854 921-1 909 924 Severity factor (S) 999 1005 1006 1004 Temperature IgT) F 939 945 946 944 Time factor( r hr 4 4.04 4 4 Pressure.' .lbs./sq. in 30 '100 100 200 Mol ratio Hilcharge (M) 2 2 2 4 catalyst fielding time hrs.. 1o 1o 1o 1o Product characteristicaz- Octane number 78. 6 8l. 5 81. 3 8l. 9 78. 6 80. 4 8 2. 81. 9 i-. '42.9 42.6 44.3 42.2 46.5 48.1 45.0 11.8 11.6 10.9 10.0

Coke .do 2.9 1.8 1.1 0.6 Difference between distillation points of charging stock and product F 18 18 25' Hydrogen produced, mols Ba/mol charge 1.1.13 1.04 0.84 1.67

. the number of mois of hydrogen per mol of Y hydrogen. A still further dift'erenatiationl is that in a destructive hydrogenation process there is inevitably a very large degradation of the charge or, in other words, the product is generally much more volatile than the charge whereas in our process the production of volatility is definitely limited and this extreme degradation is avoided.. In our process the 50% boiling point of the charge and the 50% boiling pointbf the total liquid product are within F. and usually within 50 F. of each other so that the degradation and volatility production are relatively small or even non-existent.

It is interesting to note, reverting to the matter of the ratio of' hydrogen to 'hydrocarbon charge, that the improvements inoctane number obtainable by using small concentrations of hydrogen are contra to what might logically be expected since when other conditions are held constant dilution of the charge with hydrogen results in decreasing the contact time and therefore one would expect to obtain results comparable to a lower time factor whereas actually the eii'ect of adding hydrogen within certain limits is to improve the octane number just as increasing time factor (and thereby increasing severity factor) improves it.

At the same time, as previously pointed out, 'it is important to keep the hydrogen concentration within definite, rather low limits since a point is reached at which raising the hydrogen concentration decreases instead of increasing the octane number. This is shown by Figure 7 in which a typical set of data is plotted to show this optimum and a line is shown giving the approximate locus of this optimum for different severities.

Figure 7 also illustrates the effect of mol ratio on the amount of coke (figures as in .all coke vdata given in this specification being the weight percentage based on the .charging stock). Increasing the mol ratio of hydrogen to charge markedly decreases the amount of coke. Typipally we produce from 0.5 to 3% or, more broadly, from .0.1 to 5% of coke per unit of charge.

Typical. d'ata corresponding to Figure 7- are given in the :following table to illustrate the effect of`varyinggthe mol ratio of hydrogen t..

Table 4 Example number Run number 908 921-1 860 909 Severity (S) 1006 1005 1001 1006 Temperature (T) F.. 946 945 941 940 Time factor (F) hrs 4 4. 04 4 4 Pressure lbs/sq. in., 100 100 100 100 Mol ratio Halcharge (M) 1 2 3 4 Catalyst olding time hr s 10 10 10 10 Product characteristics:

Octane number 79. 8 B1. 5, 84. 0 81. 3 Yield per cent 82.2 30. 4 79. 6 82.3 Gravity A. P. I 42. 9 42. 9 42. 4 42. 0 Aromatics per cent 44. 5 46. 5 5l. 4 48. 1 Olons-. do 11. 5 l1. 5 l0. 6 l0. 9 Coke do 2.8 1.8 1.5 1.1 Difference between 50% distillation points of charging stock and product F 17 18 2l Hydrogen pr charg 1.20 i. 04 l. 03 0. 84

As illustrated by Figure 8, the optimum mol hydrogen.

Y The area in which we nd it desirable to work is shown in Figure 8 in which mol ratio is plotted against severity factor and the cross-hatched areais the one in which we prefer to Work. Mathematically this area can be defined as lying between the severity factor limits previously given and also lying between vthe lines defined by the following equations in which M is the mol ratio and Sis the severity factor as previously dened:

Optionally the denominator in the nrst oi the above two equations can be 100 instead of 125. The preferred area lies between the preferred severity factor limits (950 and 1075) and between the preferred mol ratio limits of 0.5 and 8.

With reference to the hydrogen used, pure .100% hydrogen is theoretically best but in practice We prefer to use the impure hydrogen produced in our process which contains various amounts of methane; ethane and propane depending on the operating conditions used, the charging stock and the purification system employed. The presence of these hydrocarbons tends to result in higher coke formation but this is not serious over a considerable range. Where We speak of mol ratio of hydrogen to charge we refer to the pure hydrogen content of the gas actually used.

The examples previously given refer to the use of a molybdenum oxide on activated alumina catalyst. A catalyst supported on alumina gel gives still better results. Such a catalyst containing 3.'l5% M003 by weight can be made as follows:

Dissolve 143 kilograms of ammonium para molybdate (NH4) eMo'1O24AH2O in suiiicient distilled water to yield 2280 liters of solution. Place 1500 kilograms of granular alumina gel in an apparatus which can be evacuated and add the molybdate solution. Agitate the mixture and then attach to a source of suctionand reduce the pressure to 30-40 millimeters of mercury. Then release the vacuum and allow the pressure to rise to atmospheric. Lower the pressure a second time to the same level and then release again to atmospheric pressure. Repeat this procedure a third time and then drain the remaining liquid from the impregnated gel. Air-dry the latter is screens or other containers using layers of about 1 inch in depth.` Then place the dried material in a muille furnace in a suitable container and heat at a temperature of 1200 F. for one hour. When cooled to room temperature, the product isready to be used.

Using the catalyst thus prepared the following typical results have been obtained (Example 13) inospheric pressure.

as compared with the results obtained (Example 14) using the catalyst supported on activated alumina, the preparation of which was described earlier in this specification:

Table 5 Example No.

Severity (S) 947 V947 Temperature (T) ..F 947 947 Time factor (F) rs 1.0 1.0 Pressure s l in.. 200 I mi) Mol ratio, lig/charge (M 3. 3 3. 3 Octane number o( product 82. 0 73. 6

to 30-40 millimeters of mercury. Then release the vacuum and allow the pressure to rise to atmospheric. Lower the pressure a second time to the same level' and then release again to at- Repeat this procedure a third time and then drain the remaining liquid from the impregnated alumina. Air-dry the latter in screens or other suitable containers using layers of about 1 inch in depth. Then place the dried materialin a muiile furnace and heat it at a temperature of 1200 F. for one hour. After cooling the product is ready for use.

Using this latter catalyst the following typical results were obtained:

Table 6 Framplv number 15 Run number 1487-1 Severity factor (S). 993 Temperature (T) F 993 Time factor (F). hrs 1.0 Pressure sq. in 300 Mol ratio, Plz-containing gas/cham(` l 3. 4 Catalyst holding time hrs.. 4 Product characteristics' Octane number. 70. 4 Yield percent.. 84. i Gravity P; 49.9 Aromatics percent 34. 0 Oleilns do 12.1 Coke .dn... 0.8 Difference between 50% ion poin 0l charging stock and product F 34 1 Gas contained 49.4% Hz by volume.

. Run number i Time factor (F) All of the foregoing examples are for convenience of comparison based on charging the' 'I'his naphtha had an initial boiling point of 120 F., a 50% boiling point of 270 F., an end point of 396 F. and an octane number of 24.0. In Example Number l7-the charge was a California virgin naphtha having an A. P. I. gravity of 47.0, containing 1.5% olefins and 13.5% aromatics. This naphtha had an initial boiling point of 258 F., a 50% boiling point of 305 F., an end point of 398 F. and an octane number of 57.8. The

` catalyst used was the molybdenum oxide on activated alumina, the preparation of which is described early in this specification.

Table 7 Example number Severity factor (S) Temperature ('T) F.. hrs.. i. inv

Pressuro... Mol ratio, Hrnnntaii'iing gas/chaman Product characteristics, octane number Having thus described our process in general terms and the operating conditions which we find give us the new and unexpected results previously pointed out, we turn now to a description of specific embodiments of our invention in terms of apparatus and process now. This is illustrated in one embodiment in Figure 9 and in another embodiment in Figures 10, 11 and 12.

Figure 9 is a somewhat simplified flow diagram showing the application of our invention to a system using a stationary catalyst bed.

The feed stock which may, for instance, be a fraction containing virgin gasoline and heavier material is charged through line II by pump I2 into heater I3 from which it passes to fractionator I4 which may be o'f the usual bubble plate type. The heater and fractionator are so operated that a fraction of which at least about boils between 200 F. and 450 F. and preferably bey tween 250 F. and 425 F. and which is at least 25% and preferably at least 35% composed of paraflinic hydrocarbons having from 6 to l2 or 14 carbon atoms is Withdrawn from trapout plate I5 through valve I6 while the lighter materialJ charged through line II passes out from the top of the fractionator through line I1 and can be passedvto storage or discharged from the system through valve I8. However, this material, usual- .ly corresponding to the vaulable lighter fraction of gasoline, can be passed through valve I9 and combined with the product from our catalytic aromatization process. As will be noted, fractiontaor I4 is provided with the customary reiiux coil and can if desired be equipped with a reboiler coil as Well. Material heavier than that Vdesired as charging stock to our catalytic aroma- To illustrate the use of varied feed atleast 35% of aliphatic hydrocarbons, can come from some other source such as valved line 22, valve I6 being closed.

Examples of suitable naphtha charging stocks defined in terms of 10% and distillation points are as follows:

Whatever its source and exact nature, the charge for our process is pumped by means of pump 23 through line 24 and heat exchanger 25 where it is preheated by indirect heat exchange with the hot aromatization products and then passes through line 26 and coil 21 of heater 28 and thence to one or both of reactors 29 which are shown in the form used in a commercial embodiment of our process with catalyst tubes 30 connecting with headers 3l and 32, the space between the tubes being heated to compensate for the endothermicity of the reaction and maintain the required temperature (T). This can be accomplished by passing hot flue gas or other heating medium in through valved line 33 and out through valved line 34.A

Simultaneouslywith the introduction of the charge, hydrogen from storage tank 35 and/or recycle hydrogen from the process passes through line 36 and compressors 31 and thence through valve 38 (assuming valve 39 to be closed) and heating coil 40 and thence through valve 4I and line 42 to one or both of reactors 29 along with the heated charge which is now in the vapor phase.

Since the reaction is a dehydrogenation and aromatization reaction rather than an hydrogenation reaction and is markedly endothermic the charge and hydrogen dare generally heated to temperatures above the desired average temperature (T) in the reactors or catalyst chambers to supply the endothermic heat of reaction, and this is necessarily true if unheated reactors are used, the temperature differential depending on the process conditions and apparatus size and design. To maintain more uniform temperature and minimize any purely thermal conversion the hydrogen can be heated to a higher temperature than the charge and it can, if desired, be injected at multiple points in the catalyst bed.

Instead of using coil 21 for heating the charging stock and coil 40 for the hydrogen, valves 38, 39 and 4| can all be opened so that the charging stock and hydrogen pass together through both coils or valves 38 and 4I can be closed andvalve 39 'opened so that coil 40 is not used (or not present) and the charge and hydrogen are heated together in coil 21.

Two reactors 29 are shown. As previously mentioned, the charging stock and hydrogen are passed through one or both of them under the control of valve'slf43, 44,45 and 46. Downflow through the catalyst is illustrated and is generally preferred but upflow canlikewise be used. Also it will be apparent that a larger number of reactors can and usually will be used in order to provide a suitable regeneration cycle.

The charge and hydrogen pass through at least lery system (not shown).

sure separator 48 from which the hydrogen, usually containing some light hydrocarbon gases, passes out through line 48. All or part of this can be withdrawn from the system through valved line 50, particularly at times when it is relatively impure, and the remainder preferably passes to an hydrogenpurication system which is only shown generally since it does not constitute Aany important part of this invention. In the hydrogen purification system hydrocarbon gases and other impurities are removed by one of various methods; for instance, by scrubbing with an absorber oil. This hydrogen purification system can be by-passed by opening valve 52 and closing valve 53. Either all or part of the gases-can be sent through the by-pass' line, From the purification system and/or from the by-pass line, the hydrogen is recycled through line 36 in part but net hydrogen is withdrawn either through valved line 50, as previously described, or through pressure controlled valve 54 leading to storage tank 35.

High lpressure separator 48 can suitably be e equipped with dephlegmating coil 55.

The liquid material, including all the gasoline and heavier fractions -as well as considerable .sure than separator 48 and the gases lighter than the desired product are withdrawn from the system through pressure controlled valve 60. Part of the hydrogen. produced can also be removed at this point. This fractionator 59 is, of course, equipped with a dephlegmating coil 6| and can be equipped with a reboiler coil as well. The desired product is withdrawn through trapout plate 62 and valved line 63. The heavier material, which we sometimes refer to as' polymer and which is very limited in amount, is withdrawn from fractionator 59 through valved line4 As previously described, our process results in` considerable reduction in the deposition of coke and other carbonaceous material'as compared with prior art dehydrogenation processes operating at atmospheric pressure and/or inthe absence of hydrogen. However, it diiers from destructive hydrogenation processes in the fact that the catalyst does decrease in activity due to this deposition of coke or other carbonaceous material and it must be regenerated from time to time, usually at intervals of from to 100 hours but sometimes at shorter on-stream intervals, depending on the catalyst used, the severity factor, the mol ratio, the nature of the parainic charging stock, etc.

When regeneration is necessary or desirable using the apparatus of Figure 9, one of reactors 29 which needs regeneration ls cut out of thefjA In the meantime, of course, the remaining reactor or reactors are on stream so that the process is not interrupted.

While our invention can, as has been described,

be used in connection with a stationarycatalyst bed or a plurality of stationary catalyst beds. we

. iind that particularly good results can be obtained by the use of a moving catalyst bed and one embodiment of this type of invention is shown in Figures 10, 11 and 12. Apparatus of the Figure 10 type was used in the laboratory in obtaining most of the data given in this specification, although the data of run No. 1083A, for example, were obtained with astationary catalyst bed. l

Turning now to Figure 10 in more detail, a feed, which can have the characteristics previously described, is pumped from feed stock tank |0| by means of pump |02 through coil |03 in heater |04. Simultaneously 'hydrogen from any suitable source, including recycle hydrogen, if desired, is passed through the same heating coil (as illustrated in Figure 10) although separate heating coils can be used (as in Figure9). As shown, hydrogen from tank |05 passes through pressure reducing valve |06 and meter |01 and' through valved liney I I2 and withdrawn through valved line ||3 is preferred.

The granular catalyst of the type previously described is added by means of inlet ||r4 passing through valves H5 and I I6 separated by an injection line I|1"for. nitrogen, flue gas, or steam to act as a gas lock. The catalyst passes downward through the reaction chamber ||-0l and the hot charging stock and hydrogen pass concur'- rently therewith although counter-current ow can likewise be used. The catalyst is supported at the bottom on a rotary feeder ||8 which permits it to pass slowly downward through the reaction chamber, thence through valves I0 and |20 separated by a gas lock 2| with intermediatel nitrogen or other inert gas injection and thence to a catalyst revivication chamber which is not shown. In this revivilcation chamber the catalyst is, of course, treated with an hot oxidizing gasto burn oli thecoke and other carbonaceous material, care being taken to control the system, for instance by closing valves 43 and 45.

The systemis purged of hydrocarbons by intro-` through valved line 65, passing out of the re-' tor through valved line 66 to a suitable recov- Air dilutedwith iiue gas or some other gaseous oxidation medium is then passed through the catalyst chamber, for instance by means of valved lines 61 and 68, at controlled' temperature as known to the art, to burn of! the coke and other carbonaceous .material, after which the reactor containing the catalyst thus regenerated can again be put on stream.

temperature so that the catalyst will not be injured. 'I'he regenerated catalyst is returned to inlet II4.

Reaction products pass out from the reaction chamber in a manner which will be described in connection with Figure 12 through line |22 and thence to a trap |23, which can be packed with -glass Wool or the like, where any catalyst carried with the products is deposited and can be removed periodically. The products then go to -cooler |24 and high pressure separator |25. In

this high pressure separator the hydrogen carrying with it some light hydrocarbon gases is separated from the aromatized naphtha. The gas passes overhead through line |26 and pressure control valve |21 and thence through line |28 uable hydrocarbons. The remaining gas rich in hydrogen is vented through valved line |30. The hydrogen thus vented is substantially in excess of the' hydrogen originally introduced. Alternatively excess gas' can be removed through valved line |28a and recycled to the reaction chamber.

Turning to the liquid product from high pressure separator |25, this passes through line |3| and pressure or liquid level controlled valve |32 to a low pressure separator |33 in which some additional gas containing not only a little hydrogen and some light hydrocarbons but also an appreciable amount of butane and other valuable light naphtha constituents passes out through 'line |34, cooler |35 and line |28 to absorber |29 where it is contacted along with the gases from high pressure separator |25 with an absorption medium introduced through line |30 by means of pump |31 and withdrawn through line |38 by means of pump |39 passing through heat exexchanger |40 and heater |'4| to stripper |42 equipped with the usual reboiler coil |43 and dephlegmating coil |44. The lean oil from the stripper passes through heat exchanger |40 and cooler |45 back to pump |31 and absorber |29 for reuse. The stripper yields an overhead consisting largely of light ends of motor fuel and this passes through line |46 and cooler |41 back to low pressure separator |33. The final product is' withdrawn through valved line |48. f

Referring now to Figure 11, the granular catalyst from hopper ||4 which may be made up predominantly of granules within the range from 2 mesh to 50 mesh, for instance .4 mesh passes through valve ||5 which can be continuous or discontinuous in its operation into chamber |49 wherein inert gas is injected to prevent diffusion of naphtha vapors out of the system. Thence the catalyst passes through valve H6, which can often be omitted or left open, into a tapering tube |50 wherein the hot charge and hydrogen introduced through line |09 preheat the catalyst by indirect heat exchange and the hot catalyst passes into reaction chamber |10.

The bottom of this reaction chamber is shown at thetop of Figure 12. The catalyst passes by gravity flowi-out of reaction chamber into a .chamber |I formed by screen |52 having a mesh smaller than the smallest catalyst grains. Catalyst flow is regulated by rotary valve I8 which is operated through shaft |53 by a. variablespeed motor not shown. The catalyst then passes through valve H9, whiehcan in some cases be left open, into gaslock chamber |54 and thence Athrough continuous or intermittent valve |20 to the regeneration system. y

Having described our invention what we desire to claim is:

1. In a cyclic catalytic process including alternating on-stream and regeneration periodsfor the production 'of aromatic compounds and hydrogen from lparaiiinic hydrocarbons involving in said reaction zone of from 30 pounds to about 450 pounds per square inch gage and adding hycontacting during the on-stream period a'hydro-` carbon mixture containing at' least 35% of parainic hydrocarbons 'of from 6 to 12 carbon atoms with a dehydrogenating and cyclicizing catalyst at an elevated temperature and regulating the ow of said hydrocarbons over the catalyst with reference to the temperature maintained below the rate sufficient merelyl to convert naphthenes to aromatics and at a suiiiciently low rate largely to dehydrogenate and cyclicize said paramnic hydrocarbons to aromatics and accompanied'by the net production of hydrogen and a carbonitdrogen to said reaction zone in an amount ranging from 0.5 mol to 8 mols of hydrogen per mol of said hydrocarbon mixture and further restricted within the above range to a mol ratio determinable by the particular temperature and rate of flow of charging stock employed, the lower limit of saidmol ratio being that defined bythe equation and the upper limit by the equation Y s-soo M 15 in which S represents a severity factor defined bythe equation S=T+100 logic F in which T represents the reaction temperature employed maintained within a range not less than 875 and not more than 10'15"v F.; and F represents theftime factor being the number of hours required to flow one volume of charge` measured on the liquid basis through one volume of catalyst space and being maintained within the range between 0.1 and 25, the particular temperature and time factor employed within said ranges being related in such manner that S has a value between 925 and 1100 adapted to effect said dehydrogenation and cyclization reaction.

2. A process for converting a naphtha charging stock rich in aliphatic hydrocarbons of from 6 to 12 carbon atoms into hydrogen and a motor fuel or motor fuel component rich in aromatic hydrocarbons involving a cyclic catalytic operation including alternating .on-stream and regeneration periods which comprises contacting during the on-stream period said naphtha charging stock in the vapor phase with a dehydrogenating and cyclicizing catalyst at a temperature between 875 F. and 1075 F. maintained by supplying the endothermic heat of reaction to the catalyst contact zone, at a time factor of from 0.1 to 25, where the time factor is the number of hours required to pass one volume of charge measured as liquid through one volume of catalyst space, the time factor being selected within said range relative to the particular temperature maintained so as to effect dehydrogenation and cyclization of said aliphatic hydrocarbons and at a pressure of from 30 pounds per square inch gageto about y450 pounds per square inch gage in the presence of.

added hydrogen in an amount ranging from 0.4 mols of hydrogen per mol of said charging stock to 9 mols of hydrogen per mol of said charging stock, and removing net produced hydrogen from the process, whereby the feasible on-stream period for saidconversion is substantially increased and there is produced from said charging stock a net yield of hydrogen and a high yield of highly aromatic material boiling within the `gasoline boiling point range, having a high octane vnumber and an A. P. I. gravity less than the A. P. I. gravityof such portion of said charging stock as boils Within the gasoline boiling point 3. A process for converting a naphtha charging stock rich in aliphatic hydrocarbons into hydrogen and a motor fuel or motor iuel component rich in aromatic hydrocarbons involving a cyclic .catalytic operation including alternating on- S=T+100 logic F lying between 925 and 1100 and sumciently high to effect the conversion of aliphatic hydrocarbons to aromatic hydrocarbons, and at a pressurecf from 30 pounds per square inch gage to about 450 pounds per square inch gage in the presence of added hydrogen in an amount ranging from 0.4 mois of hydrogen per mol of said charging stock to 9 mols of' hydrogen per mol of said chargequation S- 875 M 125 l and not more than that defined by the equation S -890 ing stock, the mol ratio of hydrogen to charging stock being selected on the basis of the severity factor in such manner that the mol ratio is not less than that deined bythe equation and not more than that dened by the equation s-sco M 15 and removing net produced hydrogen from the process, whereby the feasible on-stream period for said conversion is substantially increased and carbons present in a naphtha charging stock into' hydrogen and-aromatic hydrocarbons involving a cyclic catalytic operation including alternatingori-stream and regeneration periods which comprises flowing during the on-stream period said naphtha charging stock in the vapor phase at a temperature (T) between 875 F. and 1025 F. through areaction zone in contact withl a dehydrogenating and cyclicizing catalyst comprising an oxide of a metal selected from the left hand columns of groups IV, V and VI of the periodic table, maintaining said temperature by supplying the endothermic heat of reaction tov the catalyst contact zone, regulating the rate' of ilow of said naphtha charging stock through said dehydrogenating and cyclicizing catalyst to provide a time factor between 0.1 and 25, where time factor (F) is the number o'f hours required to now one volume of charge measured as liquid through one volume of catalyst space, maintaining the severity factor (S) defined by the equation between 925 and 1075 and sumciently high to e'ect the conversion of aliphatic hydrocarbons and removing net produced hydrogen from the I process, whereby the feasible ori-stream periodA for said conversion is substantially increased and said aliphatic hydrocarbons are converted in large measure into aromatic hydrocarbonsboiling within the gasoline boiling point range and having high octane numbers and whereby a net yield of hydrogen and a product boiling within the gasoline boiling-point range having an A. P. I. gravity less than the A. P. I. gravity of such portion of said charging stock as boils within the gasoline boiling point'range are produced.

5. A process for converting a naphtha charging stock containing a substantial amount of aliphatic hydrocarbons into hydrogen and a motor iuel or motor fuel component rich in aromatic hydrocarbons involving a cyclic catalytic operation including alternating ori-stream and regeneration periods which comprises contacting during the ori-stream period said naphtha charging stock in the vapor phase in a'reaction zone :in the presence of added hydrogen with adehydrogenating and cyciicizing catalyst comprising an oxide of a metal selected fromlthe le'ft hand column of group VI supported on alumina at a temperature (T) between 800 Rand 1025 F., supplying heat to said reaction zone. in addition to that necessary to heat said charging stock and said added hydrogen to the desired temperature (T) thus supplying the endothermic heat of reaction, at a time factor of from 0.2 to 20.

`where time factor (F) is the number ofhoursf required to contact one volume of charge measured as liquid through one volume of catalyst space', at a severity factor (S) dened by the equation i Sv=T+ logw F lying between 975 and 1075 adapted to convert the mol'ratio is not iess-than that defined by the equation and not more than that defined by the equation and removing net produced hydrogen from the process, whereby the feasible on-stream period for said conversion is substantially increased and said aliphatichydrocarbons are converted in large measure into aromatic hydrocarbons boiling within the gasoline boiling po'int Arange and having hi'gh octane numbers and whereby a net yield of hydrogen and a'product boiling within the gasoline boiling point range having an A. P. I. gravity less than the A. P. I. gravity of said charging stock are produced.

6. A process for converting a naphtha charging `stock containing a substantial amount of aliphatic hydrocarbons into hydrogen and a motor fuel or motor fuel component rich in aromatic hydrocarbons involving a cyclic catalytic operation including alternating on-strearn and regeneration periods which comprises. contacting said naphtha charging stock in the vapor phase in a reaction zone in the presence of'added hydrogen with a dehydrogenating and cyclicizing catalyst comprising --molybdenum oxide supported on alumina at a temperature (T) between 890 F. and 1025 F., supplying heat to said reactionvzone in addition to that necessary to heat said charging stock and said added hydrogen to the desired temperature (T) thus supplying the endothermic heat of reaction, at a time factor of from 0.2 to 20, where time factor (F) is the number of hours required to contact one volume of charge meas- ..ured as liquid through one volume of catalyst space, at-a severity factor (S) defined by the equation S=T+100 logro F lying between 975v and 1075 adapted-to convert aliphatic hydrocarbons to aromatic hydrocarbons, and at a pressure of from 50 pounds per 4square inch .gage to 300 pounds per square inch gage in the presence of hydrogen added in an .amount ranging. from 0.5 mol offhydrogen per mol of said charging stock to 8 mols of hydrogen pei-mol of said charging stock, the mol ratio of *hydrogen to charging stock being selected, on the basis of the severity factor in'such manner that the mol ratio is not less than thatdefined by the equation a s- 875 Mft-125 and not more than that defined by the equation i s-soo M- cyclicizing catalyst consisting essentially of molybdenum oxide adsorbed on alumina, main-v taining said temperature by supplying the endothermic heat of reaction to the catalyst contact zone regulating the rate of flow of said charging stock through said moving bed of catalyst to provvide a time factor'between 0.1 and 25, where time factor (F) is the number of hours required to contact one volume of charging stock measured as liquid through one volume of catalyst space, maintaining the severity factor (S) dened by the equation u S=T+1oo log F the reaction zone in an amount ranging from 0.5 mol of hydrogen per mol of said charging'stock to 9 mols'of hydrogen per mol of said charging stock, the mol ratio of hydrogen to charging stock being selected within this range on the basis of the severity factor in such manner that the procesa whereby the feasible' n-stream period for said conversion is substantially increased and said aliphatic hydrocarbons Yare converted in large measure vinto aromatic hydrocarbons boiling within the gasoline boiling point range and ,having high octane number ,and whereby a net vyield of hydrogen and a productlboiling within the gasoline boiling point range having an A. P.v I. gravity at least 3 A. P. I. less than the A. P. I. vgravity of that part of said charging stock which boils within the gasoline `boiling point range are produced.

7. A process for converting a naphtha charging stock containing a large amount of aliphatic hydrocarbons into hydrogen and an aromatic motor fuel or motor fuel component involving a cyclic catalytic operation including alternating mol ratio is not 'less than that dened by the equation and not more than that defined by the equation s-sso M =1"5 and removing net produced hydrogen from the process whereby the feasible on-stream period for said conversion isvsubstantially increased and aliphatic hydrocarbons present in said charging n stock are converted in large measure into aromatic hydrocarbons boiling within the gasoline boiling point range and having high octane numbers and whereby a net yield of hydrogen and a Y productvboiling within the gasoline boiling point .range having an A. P. I. gravity less than the A. P. I. gravity of such portion of said charging stock as boils within the gasoline boiling point range are produced.

8. Avprocess according to claim 7 in which said molybdenum oxide constitutes from 2% to `10% by weight of the total catalyst including the support.

9. A process for converting a naphtha charging stock rich in aliphatic hydrocarbons into hye drogen and a motor fuel or motor fuel component rich-in aromatic hydrocarbons involving a cyclic catalytic operation including alternating onstream and regeneration periods which comprises contacting-during the cn-stream period said aliphatic naphtha.' chiarging stock in the vapor phase with a dehydrogenatin'g and cyclicizing catalyst at a temperature (T) between 875 F. and 1075 F., maintained by supplying the endothermic heat of reaction to the catalyst contact zone, at a time factor of from 0.1'to 25. where the time factor (F) is the number of hours required on-stream and regeneration periods which com-,1.

" `prises owing durirg the on-stream period said ,naphtha charging stock in the vapor phase at a temperature (T) between 875 F. and 1025" F.

through a moving bed of' a dehydrogenating" and to pass one volume of charge. measured as liquid through one volume of catalyst space, at a severity factor (S) defined by the equation lying between -925 and 1100 and sufficiently high to eectl the conversion of aliphatic hydrocarbons to aromatic hydrocarbons.- and at a pressure of from 30 pounds per square inch gage to about 450 pounds per square inch gage in the presence of added hydrogen in an amount rang-l ing from 0.4 mol of hydrogen per mol of said charging stock to 9 mois of hydrogen per moi of said charging stock, the mol ratio of hydrogen to charging stock being selected on the basis of the severity factor in such manner that the mol ratio is not less than that defined by the equation s 10o and not more than that defined by the equation s-soo M" and removing net produced hydrogen from the 4 10. A process for converting a naphtha charg" ing stock containing a large amount of paramnic hydrocarbons into hydrogen and a motor fuel or motor fuel component rich in aromatic hydrocarbons involving a. cyclic catalytic operation including alternating ori-stream and regeneration periods Whichcomprises contacting during the on-streani period said paralnic naphtha'charging stoel: in the vapor phase with a dehydrogenating and cyclicizing catalyst at a temperature (T) between--875 F'. and 1025 F., maintained by supplying the endothermic heat of reaction to the catalyst contact zone, at 'a time factor of from 0.1 to 25, where the time factor (F) is the number of hours required to pass one volume of charge measured as liquid through one volume of catalyst space, at a severity factor (S) defined by the equation s=r+1co logro F l lying between 925 and i075 and suiclently high to elect the conversion of aliphatic lwdrocarbons to aromatic hydrocarbons. and at a pressure of from pounds per square inch gage to about 450 pounds per square inch gage in the presence of added hydrogen in an `amount ranging from. 0.5 mol of hydrogen per moi of said charging stock to 9 mois of hydrogen per moi of said charging stock, the mol ratio oi hydrogen to charging stock being selected on the basis oi the severity factor in such manner that the mol ratio is not less4 than thatideflned by the equation ^ a dehydrogenating and cyclicizing catalyst -comwithin the gasoline boiling point range, having a high octane number and an A. P. I. gravity less than the A. P. I. gravity of such portion of. said charging stock as boils within the gasoline boiling point range, and whereby said catalyst is fouled by coke-like carbonaceous matter at' a rate much less than that for ordinary dehydrogenation processes but much greater' than that for hydrogenation processes, and following lthe on-stream period regenerating said catalyst'by contacting it with a hot oxygen-containing gas.

1l. A process for converting a naphtha charging stock containing at least 35% of paraiiinic hydrocarbons into hydrogen and a motor -fuel or motor fuel component containing at least v35% of aromatic hydrocarbons involving a cyclic catalytic operation including alternating ori-stream and regeneration periods which comprises ccntacting during the on-stream vperiod said naphtha charging stock in the vapor phase at a temperature (T) between 875 F. and 1025. F., with prising an oxide of a metal selected from the left hand columns of groups IV, V and VI of the periodic table, maintaining said temperature by applying the endothermic heat of retaction to he catalyst contact zone, regulating the rate of ow of said charging stock through said catalyst to provide a timefaetor between 0.1 and '25, where time factor (F) is the number of hours required tocontact one volume of charge measured as liquid. through one volume of catalyst space, maintaining the severity factor. (S) defined by the equation S=T+100 logro F between 925 and 1075 and suiiiciently high to effect the conversion of`-aliphatic hydrocarbons to aromatic hydrocarbons, maintaining the pres-l sure of said contacting operation at from 30 pounds per square inch gage to about 450 pounds per square inch gage,l introducing hydr' and not more than that defined by the equation S-sgo Mf- 15 recycling a part of the hydrogen produced by u the reaction to the reaction zone and removing net produced hydrogen from the process, whereby the feasible ori-stream period for said conver- -sion is substantially increased and said paraiiinic 1 hydrocarbons are converted in large measure into aromatic hydrocarbons boiling within the gasoline boiling point range and having high octane numbers and whereby' a net yield of hydrogen and a product boiling within the gasoline boiling point range having an A, IJ. I. gravity less than the A. P. I. gravity of such portion of said charging stock as boils within the gasoline boiling point range are produced. l

l2. A. process as defined in claim 9 wherein 'said naphtha charging stock comprises a cracked naphtha and wherein said pressure is maintained a 2,320,147" I in the range of about 30 to 250 pounds per square inch gauge. A1

13. A process as defined in claim 9 vwherein said charging stock consists essentially of a virgin petroleum' naphtha and said pressure is maintained at a minimum value in excess of 100 lbs. per square inch gauge.

14. A cyclic catalytic process of converting a the net production of hydrogen and the Adeposition of a deactivating carbonaceous material on the catalyst, said added free hydrogen being introduced in amount suillcient to maintain a feed ratio of hydrogen to naphtha of about 0.5 to `8.0 mols per mol of naphtha, and maintaining the combined partial pressures of said hydrogen and naphtha in thereaction zone within the range of from 30 to about 450 pounds per. square inch gauge, Vthereby' eecting an enhanced yield of Aaromatic compounds and a substantial decrease aceous material on the surface of the catalyst produced by side reactions incident to said conversion, and thereafter during a regeneration period, the activity of said catalyst for said conversion is revivied, and said on-stream conversion and catalyst revivincation periods are alternately and continuously repeated in a cyclic operation, the improved method whereby an enhanced yield 0f the desired aromatic products is produced, comprising adding free lwdrogen to said stream of aliphatic hydrocarbons in an amount not less than about 1/2 mol volume and not in excess of about 8 mol volumes per mol volume of hydrocarbon feed whereby a concentration of free hydrogen in substantial excess of the free hydrogen concentration resulting from the conversion alone is maintained in the catalytic contacting zone, and maintaining said zone under a pressure of not less than about 30 and not in excess of about 450 lbs. per square inch gauge.`

in the quantity of laid carbonaceous deposit produced per unitof naphtha charged and a correspending increase in the feasible :on-stream period before regeneration of the catalyst is required compared With a similar operation except in the absence of said added hydrogen and maintained pressure, continuing the passage of said naphtha vapors and hydrogen over the catalyst until the carbonaceous deposit thereon accumulates to an extent suiilcient to decrease the conversion .of aliphatic compounds to hydrogen and aromatics to an undesired extent, and thereafter regenerating the catalyst for reuse in said conversion by removal of the carbonaceous deposit therefrom.

15. In a process for producing hydrogen and aromatic hydrocarbons from aliphatic hydrocarbons having boiling points within the boil- .ing point range of naphtha wherein, during an on-stream period, a stream containing aliphatic hydrocarbons is passed in contact with a de'- hydrogenating and cyclicizing catalyst under conditions adapted to convert said hydrocarbons to hydrogen and aromatic hydrocarbons and the flow of said hydrocarbons in contact with the catalyst vis continued until the convyersion to aromatica and hydrogen decreases to an undesired extent due to an accumulation of carbon-` 16. In a process for' producing hydrogen and aromatic hydrocarbons from aliphatic hydrocarbons having boiling points within the boiling point range of naphtha wherein, during an on- -stream period, a stream containing aliphatic hydrocarbons is passed in contact with a dehydrogenating and cyclicizing catalyst under conditions adapted to convert said hydrocarbons primarily to hydrogen and aromatic hydrocarbons with incidental production due to side reactions of appreciable but limited quantities of olefinic normally liquid hydrocarbons and cracked normally gaseous hydrocarbons and the flow continued of said hydrocarbons in contact with the vcatalyst until the conversion to aromatics and hydrogen decreases to'an undesired extent due to an accumulation of solid carbonaceous material on the surface of the catalyst also produced by said side reactions, and thereafter, during a regeneration period, the activity of said catalyst for said conversion is reviviiled.- and said on-stream conversion and catalyst revivication periods are alternately and continuously repeated in .a cyclic operatiomthe improved method whereby an enhanced yield ofthe desired aromatic products is produced, comprising adding free hydrogen to said stream of aliphatic hydrocarbons in an amount not less than about 1/2' mol volume and not in excess of about 8 mol volumes per mol volume of hydrocarbon :feed whereby a concentration oi free hydrogen in -substantial excess of the free hydrogen concentration resulting from the conversion alone is maintained in the catalytic contacting zone, and maintaining said zoneunder a pressure of not less than about 30 and not in excess of about 450 lbs. per square inch gauge.

EDWIN T. LAYNG.

LOUIS C. RUBIN.

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