Process for preparing aromatics from naphtha

This invention is directed to producing a high yield of aromatics from naphtha by integrating a catalytic reformer and a thermal hydrocracking unit followed by fractionating the product stream to obtain the desired aromatic. In this process, after sulfur is removed, the naphtha is catalytically reformed under conditions adapted to provide the maximum aromatic content. The reformed stream is then thermally hydrocracked, preferably in a system containing multiple reactors. The process does not require an external source of hydrogen since the reforming and hydrocracking units are operated to maintain a hydrogen balance. The hydrocracked stream is fractionated in a series of columns to produce the desired mix of aromatics. In one example, the fractionation separates out (1) benzene, which is further purified, if necessary (2) a toluene stream and a C.sub.9 + aromatic stream which are charged to a transalkylation reactor to produce a composition containing mixed xylenes sent to recycle and (3) a mixed xylene stream from which paraxylene is separated by a liquid-solid adsorption technique and from which the remaining xylenes are isomerized to produce further paraxylene. A by-product stream of light hydrocarbons is also produced. This procedure can be readily modified to produce greater amounts of benzene and/or toluene and less xylene.

BACKGROUND OF THE INVENTION 
In the reforming of naphtha to produce high quality gasoline substantial 
amounts of mixed aromatics are produced, particularly 
benzene-toluene-xylenes and ethylbenzene. An approach which has been tried 
for the production of aromatics, primarily a substantially pure stream of 
benzene from naphtha, is set forth in British Pat. No. 1,157,114. This 
patent discloses the integration of a reforming unit and a thermal 
hydrocracking unit for producing predominantly benzene and a fuel gas from 
naphtha. In a first stage endothermic reaction, a naphtha and hydrogen 
stream are reacted at about 500 psig in a catalytic reforming operation to 
produce aromatic compounds from paraffins, olefins and cyclic 
hydrocarbons. In a second stage exothermic reaction at about the same 
pressure, the reformate stream is subjected to thermal hydrocracking to 
produce methane and ethane from paraffins, and to dealkylate aromatic 
compounds. Steam and/or sulfur is added to prevent coking in the thermal 
hydrogenation reactor. The principal product stream of this process is a 
mixture of about 75% benzene, 14% toluene and 7.5% xylene which is then 
further treated to produce substantially pure benzene, a toluene-rich 
stream of benzene-toluene-xylene, and a xylene-rich stream containing 
higher aromatics. The second major product stream is a fuel gas rich in 
methane, ethane, hydrogen, carbon monoxide and carbon dioxide, and 
containing sulfur compounds. Although a hydrogen recycle stream may be 
used, an outside source of hydrogen is required to carry out the processes 
of this patent. 
It is an object of this invention to provide an improved process for 
producing aromatics from naphtha. For example, it would be desirable to 
obtain a more efficient economical conversion of naphtha to aromatics such 
as benzene, toluene, and xylenes and to do so without an outside source of 
hydrogen. 
It also would be desirable to upgrade in value the products of the 
foregoing process by increasing the xylene, and particularly the 
paraxylene content for use as a feedstock for the production of 
terephthalic acid or dimethylterephthalate. Furthermore, it is desirable 
to increase the available ethane, which as a feedstock for the production 
of ethylene has a higher value than methane. 
Of the normal amount of mixed benzene-toluene-xylenes production, the minor 
proportion of the xylenes are recovered, for use as solvents and for 
chemical uses such as the production of terephthalic acid from paraxylene. 
The C.sub.8 aromatics found in catalytic reformate generally occur in the 
following ratio 
______________________________________ 
Typical % Range % 
______________________________________ 
Orthoxylene 23 23-26 
Metaxylene 40 35-40 
Paraxylene 17 16-20 
Ethylbenzene 20 17-21 
______________________________________ 
The total amount of these aromatics in the reformate can be about 15 to 
18%, although the amount can vary significantly depending upon the 
character of the feedstock and the processing conditions. However, even 
when a catalytic reformer is operated under conditions to maximize the 
production of benzene-toluene-xylenes further substantial efforts and 
additional costs are necessary to separate the components of this mixture 
and to produce a high-purity paraxylene. The paraxylene separation is 
usually carried out by crystallization or adsorption and the higher the 
content of paraxylene, the higher the value of the C.sub.8 aromatic 
fraction. It is therefore preferable to seek first to increase the ratio 
and amount of xylenes in the product stream by chemical processing steps. 
It is therefore a further object of this invention to produce a substantial 
yield of high purity paraxylene from a relatively wide cut, low quality 
naphtha. In addition, it is desirable to produce in this process nitration 
grade benzene, pipe line quality gas high in ethane, propane, and small 
quantities of higher boiling paraffin compounds. 
Another object of this invention is to carry out hydrodealkylation of the 
effluent from a catalytic reformer operated on high end point naphthas so 
that the effluent contains C.sub.9, C.sub.10 and C.sub.11 aromatics, 
making it possible to: 
increase the total aromatic content of the effluent to produce and recover 
as much as 62%, or more, aromatics from an Arabian naphtha; 
minimize the formation of methane in the gas produced, and maximizing the 
ethane content; 
hydrocrack substantially all of the paraffins in the effluent to methane, 
ethane, and propane without the need for hydrogen beyond that produced in 
the reformer; 
produce a mixed xylene fraction with an enhanced paraxylene content; and 
accomplish the foregoing through processing techniques that reduce 
investment and operating costs substantially below conventional methods. 
Other objects will be set forth in the following detailed description of 
the invention. 
SUMMARY OF THE INVENTION 
This invention is directed to a process for treating naphtha to produce a 
substantial amount of aromatic compounds such as benzene, toluene and 
xylene. The invention will be explained primarily by setting forth a 
process for the production of paraxylene and by-product streams of 
benzene, propane and fuel gas. As set forth below, this process can be 
readily modified to make greater quantities of benzene and/or toluene with 
concommitantly lower amounts of xylenes. 
The process steps for making paraxylene include naphtha pretreatment, 
catalytic reforming, thermal hydrocracking, gas recovery, aromatics 
fractionation, aromatics transalkylation, paraxylene recovery and xylenes 
isomerization. A heavy naphtha, having a boiling point range of about 
212.degree. F to 374.degree. F, is initially desulfurized by hydrogen 
treatment and other impurities are removed to prevent catalyst poisoning 
in subsequent steps. 
The treated product stream is passed through a series of reforming reactors 
and then directly to a thermal hydrocracking reactor. These reactors are 
substantially in hydrogen balance so that an external source of hydrogen 
is unnecessary. The reformer is operated at a low-pressure and high 
severity to promote the formation of aromatics from paraffins. The thermal 
hydrocracking reactor is operated to convert most of the heavier paraffins 
to lighter paraffins and the heavy aromatics are partially demethylated. 
The rate of the hydrocracking or dealkylation reaction increases with the 
molecular weight of the aromatics so that this reaction can be controlled 
to substantially increase the concentration of toluene and xylene. 
The hydrocracking reactor outlet is cooled and separated into liquid and 
gas streams. By-product hydrogen and other gases are recovered by 
treatment through a compressor, absorber and cryogenic unit. The hydrogen 
is used in the several recycle streams in the process, primarily in the 
reformer and hydrocracking unit. The other gases which are separated may 
be used for fuel or subjected to further separation techniques for 
conversion to other products. The liquid product stream is optionally 
treated in a stabilizer to remove propane and other gases. 
The major product stream then passes through a deisohexanizer from which 
paraffins are recovered overhead. The bottoms product stream and recycle 
production from downstream operations are passed through a series of 
distillation columns to separate the aromatic compounds. The first tower 
produces overhead a benzene stream which may contain some paraffins and 
other compounds. Depending on the amount, if any, of paraffins the benzene 
may be charged to an extractive distillation tower which recovers pure 
benzene and a by-product paraffin stream which may be used for fuel. The 
bottoms from this benzene tower are charged to the next tower from which 
toluene is removed overhead for use downstream in a transalkylation unit. 
The toluene tower bottoms are charged to the next tower from which mixed 
xylenes are removed overhead and sent to a paraxylene recovery operation. 
The xylene tower bottoms go to a rerun column from which an overhead 
C.sub.9.sup.+ stream is recovered for use in the transalkylation unit. The 
rerun column bottoms are heavy aromatics which may be used as fuel or 
partially recycled. 
Alternatively, the xylene tower bottoms may be recycled to the thermal 
hydrocracking reactor and the transalkylation process eliminated from the 
processing sequence. The toluene bottoms can be used as a final product or 
converted to benzene -- if additional xylenes are not desired. 
The transalkylation unit produces mixed xylenes from a mixture of higher 
and lower aromatics. The overhead from the rerun column and toluene column 
and some of the bottoms from the xylene column are charged to the 
transalkylation unit. A catalytic transfer of methyl groups occurs to 
produce an equilibrium mixture of xylenes, benzene, and some heavier 
aromatics. The liquid product is recycled to the aromatics fractionation 
section and the gas output can be used for fuel. 
A paraxylene adsorption recovery unit is used to separate paraxylene from 
the mixed xylenes obtained from the overhead of the xylene column. 
Paraxylene is recovered by adsorption from the liquid phase in a fixed bed 
of solid adsorbent. The alternative method commercially used is a 
low-temperature crystallization process to separate paraxylene. 
A xylene isomerization unit is used to prepare an equilibrium mixture of 
xylenes from the remaining by-product xylenes of the paraxylene recovery 
unit. These remaining xylenes (mainly metaxylene) are catalytically 
isomerized to a mixture of orthoxylene, metaxylene, and paraxylene, which 
is then recycled to the aromatics fractionation columns. Paraxylene 
comprises about 25% of the mixed xylenes obtained by isomerization. 
The invention will be explained in greater detail in connection with a 
description of the attached Figures, and the embodiment of a typical run, 
based on computer derived standard operating conditions.

THE INVENTION 
Arabian light crude oil may be chosen as the feedstock for this process. A 
large volume of it is readily available, its properties are well known, 
and naphthas cut from it are less suitable for other uses because of 
relatively poor catalytic reforming characteristics. Of course, naphthas 
from other crude oils may also be used. 
The boiling range of the naphtha to be processed can be chosen from many 
available cuts. A wide boiling range naphtha (e.g. 160.degree. to 
400.degree. F) may be used in this invention as opposed to conventional 
low end point narrow boiling range naptha used for benzene-toluene-xylene 
operations which restricts feedstock availability and increases naptha 
costs. Since hexanes and heptanes are difficult to convert and contribute 
little to the yield of paraxylene it is advantageous to substantially 
eliminate this fraction from the front end. The end point can be chosen to 
eliminate some C.sub.12 and heavier aromatics in order to simplify side 
effects. A naphtha boiling range of 212.degree. to 374.degree. F is a 
preferred range. 
Table I sets forth a typical set of properties, and is the charge stock 
used for the embodiment of this invention as set forth in the series of 
tables which follows. 
TABLE I 
______________________________________ 
CHARGE STOCK PROPERTIES 
.degree. API - 56.9 
Boiling Range .degree. F 212/374 
Crude Source: Arabian Light 
Yield Cut Range Volume % Crude: 11.2 - 27.6 
Yield on Crude Volume % 16.4 
Total Sulfur, Weight % .032 
FEEDSTOCK COMPONENTS 
Naphthenes Vol. % 
C.sub.6 0.6 
C.sub.7 2.1 
C.sub.8 3.7 
C.sub.9 4.4 
C.sub.10 4.6 
C.sub.11 4.7 
TOTAL 20.1 
Aromatics 
Toluene 1.8 
Ethyl Benzene 0.6 
O-Xylene 1.7 
M & P Xylene 2.2 
C.sub.8 Aromatics 4.5 
C.sub.9 Aromatics 4.4 
C.sub.10 4.4 
C.sub.11 1.6 
TOTAL 16.7 
Paraffins 63.2 
TOTAL 100.0 
Charge: 20,000 BPSD 
______________________________________ 
The naphtha is first mixed with a hydrogen-rich recycle stream, in a mol 
ratio of about 1:2 - 7 under a pressure of about 400 to 600 psi, preheated 
to about 650.degree.-750.degree. F, and passed over a conventional 
desulfurization catalyst (i.e. cobalt molybdenum or nickel molybdenum) in 
reactor 1 for the removal of sulfur and other contaminants which may be 
injurious to the catalytic reforming catalyst. The treated naphtha is 
cooled to about 120.degree. F, recycle gas and liquid separated and the 
light ends (such as butane and lighter) are removed in a stripper tower 2. 
A tail gas containing the hydrogen sulfide is extracted in the stripper. 
The tail gas is treated in a conventional manner (not shown) to remove the 
H.sub.2 S in an amine unit after which the H.sub.2 S is then converted to 
elemental sulfur and the tail gas is treated in a hydrogen recovery unit 
for use in the recycle stream. 
The treated naphtha is mixed with a hydrogen-rich recycle gas, preheated 
and charged to a semi-regenerative type reformer 3 (one or more stages, 
preferably three stages with interstage heating -- not shown), operated at 
severe conditions chosen to give a six months cycle. The operation of this 
unit is adapted to ensure the information and retention of the maximum 
number of aromatics rings. Table II sets forth typical conditions for a 
`Rhenforming` catalytic reforming useful in this invention. 
TABLE II 
______________________________________ 
Process Conditions Ex. I Ex. II 
______________________________________ 
LHSV hr .sup.-1 0.88 1.06 
H.sub.2 /HC mol/mol 
6.0 6.0 
PSIG-Reactor outlet 
200 275 
C.sub.5 + Octane 102 102 
Run length-months 51/2 6.0 
.degree. F (Start of Run) 
955 970 
Max .degree. F (End of Run) 
1025 1025 
______________________________________ 
Rhenforming is one of a number of commercially used catalytic reforming 
processes; it employs platinum-ruthenium catalyst which was the basis for 
the data of Table II in which the catalyst was Type E (see The Oil and Gas 
Journal, pp. 121-130 (May 1, 1976)). This process may be carried out at 
pressures from about 125 to 275 psig and at average reaction temperatures 
of about 900.degree. to 1100.degree. F, preferably about 955.degree. to 
1025.degree. F, with a liquid hourly space velocity of about 0.7 to 3.1, 
preferably about 0.8 to 1.1 and a molar ratio of hydrogen to hydrocarbons 
of about 3 to 11, preferably about 5 to 7. 
Table III sets forth typical reformer yield estimates for the process of 
this invention. In this and subsequent tables, the prefix N is for normal 
(or straight-chain hydrocarbons), and I is for iso (or branched-chain 
hydrocarbons). The suffix P is for paraffin, N is for naphthene and A is 
for aromatic. The abbreviation Bz is for benzene, T for toluene, EB for 
ethyl benzene, Px for paraxylene, Ox for orthoxylene and Mx for 
metaxylene. The data is based upon the charge stock of Table I. 
TABLE III 
______________________________________ 
CATALYTIC REFORMER NET PRODUCT YIELDS 
Vol. % B/D Wt. % #/Hr. Mols/Hr. 
______________________________________ 
H.sub.2 2.89 6,325 3,137.4 
C.sub.1 1.19 2,604 162.7 
C.sub.2 2.23 4,880 162.1 
C.sub.3 3.00 6,566 148.8 
IC.sub.4 
1.94 388 1.44 3,152 54.3 
NC.sub.4 
2.81 562 2.17 4,749 81.7 
IC.sub.5 
3.00 600 2.47 5,406 75.0 
NC.sub.5 
1.98 396 1.65 3,611 50.0 
C6P 6.37 1,275 12,349 143.3 
C7P 3.97 795 7,999 79.8 
C8P 1.00 200 2,039 17.9 
C9P+ 1.40 280 2,659 20.8 
C6N .10 20 218 2.6 
C7N .40 80 932 9.5 
C8N+ .20 40 474 4.2 
Bz .50 100 1,288 16.5 
T 6.20 1,240 15,754 171.0 
EB 3.47 694 8,817 83.0 
px 4.16 832 10,498 99.0 
m-x 9.70 1,940 24,580 231.4 
o-x 5.78 1,156 14,910 140.4 
C.sub.9 A 
20.15 4,030 51,418 427.8 
C.sub.10 A 
9.30 1,860 23,471 174.9 
C.sub.11 A 
1.55 310 3,858 26.0 
TOTAL 83.99 16,798 218,557 
5,520.1 
______________________________________ 
The effluent from the last reactor of the catalytic reforming unit passes 
directly through a heat exchanger 6 and then through start-up heater 4 and 
into a thermal hydrocracking reactor 5 which operates at a pressure range 
between 125 to 275, or up to 375 psig. In the illustrated procedure the 
pressure was 180 psig and the outlet temperature was maintained below 
about 1350.degree. F. The reaction in this reactor is exothermic but 
carefully controlled by mixing with feed and recycle hydrogen which 
normally eliminate the need for any other quench. The residence time in 
the thermal hydrocracking unit is short and with the low pressure and 
temperature the degree of reaction is held to a level where a heat balance 
is maintained. This operation is adapted to convert aromatic compounds 
containing 9, 10, and 11 carbon atoms into benzene, toluene, and xylenes 
thereby: 
1. increasing the total xylenes produced from the naphtha feed; 
2. adjusting the ratio of toluene to C.sub.9 plus aromatics to nearly an 
equal molar ratio so that maximum xylenes may be produced by 
transalkylation; 
3. substantially reducing or eliminating all paraffins boiling in the 
xylene boiling range. This reduces the additional demands that would 
otherwise be placed on the xylene isomerization catalyst and reduces the 
size of the paraxylene extraction unit and xylene isomerization units and 
reduces their operating cost. This approach is more economical than 
solvent extraction of the C.sub.8 fraction which also leads to loss of 
some aromatics to the raffinate; 
4. converting high boiling paraffins into methane and ethane and propane 
which have a higher monetary value per unit heating value; 
5. reducing the paraffin content of the benzene cut thereby making possible 
the production of benzene of nitration grade by fractionation or by 
extractive distillation at lower cost than by liquid/liquid solvent 
extraction; 
6. converting some paraffins and isoparaffins into aromatics at the 
operating pressure proposed. Process yields presented below do not take 
this reaction into account but would improve the yield by several percent; 
and 
7. converting indan into benzene and light hydrocarbons. While only about 
2% of the C.sub.9 fraction from catalytic reforming is indan, it is a 
catalyst poison to the transalkylation reaction. Although modest amounts 
of indan may be tolerated in the combined feed to the transalkylation 
unit, some compensation in operating conditions is required such as 
temperature compensation to offset depression of catalyst activity. The 
C.sub.9 cut is usually fractionated to remove most indan. This requires 
rejection of some trimethylbenzene to the indan fraction, ultimately 
reducing paraxylene yields, and also requires high investment in a C.sub.9 
fraction splitter and associated high operating costs. By use of the 
thermal hydrogenation reactor, indan is reduced to levels such that a 
splitter is not required. The transalkylation unit may still require a 
rerun column to remove heavy aromatics formed during the reaction but this 
is a much smaller less precise column. 
Table IV sets forth typical thermal hydrocracking yield estimates for the 
process of this invention. 
TABLE IV 
______________________________________ 
THERMAL HYDROCRACKING 
Mols/Hr. Net Out 
Reactor In Out Recycle 
Net Out 
Lbs./Hr 
______________________________________ 
H.sub.2 13,238.4 11,538.7 10,101.0 
1,437.7 
2,898 
C.sub.1 1,757.7 3,149.5 1,595.0 
1,554.5 
24,872 
C.sub.2 1,757.1 2,192.9 1,595.0 
597.9 17,997 
C.sub.3 348.8 497.1 200.0 297.1 13,102 
C.sub.4 136.0 174.3 174.3 10,127 
C.sub.5 125.0 61.0 61.0 4,398 
C.sub.6 
Saturates 
145.9 22.0 22.0 1,896 
C.sub.7 
Saturates 
89.3 4.8 4.8 481 
C.sub.8 
Saturates 
22.1 0.7 0.7 80 
C.sub.9 
Saturates 
20.8 0.2 0.2 26 
Benzene 16.5 67.7 67.7 5,287 
Toluene 171.0 350.3 350.3 32,260 
C.sub.8 A 
553.7 671.6 
EB 100.7 10,694 
px 120.1 12,755 
mx 280.6 29,800 
ox 170.2 18,075 
C.sub.9 A 
427.8 272.4 272.4 32,742 
C.sub.10 A 
174.9 7.4 7.4 993 
C.sub.11 A 
26.0 0.5 0.5 74 
TOTAL 5,520.1 
218,557 
______________________________________ 
The conventional thermal hydrocracking operation in commercial operation is 
the hydrodealkylation of toluene to benzene and methane. The toluene is 
produced by using a solvent to extract aromatics from a fractionated 
reformate and by then fractionating the extracted aromatics to obtain 
toluene. Toluene is a relatively stable compound and requires severe 
conditions for the reaction to proceed and to obtain a reasonably high 
conversion per pass to benzene. For example to convert 75% of toluene per 
pass through the reactor, the conditions in a commercial unit may be: 
______________________________________ 
Reactor pressure 600 psig 
Inlet temperature 1,200.degree. F 
Outlet temperature 1,360.degree. F 
H.sub.2 /toluene at inlet 
7/1 mol/mol 
Reactor residence time 
25 seconds 
______________________________________ 
Time and temperature are to some degree interchangeable, with lower 
temperatures and longer time giving equal conversions. A high hydrogen 
pressure is necessary to suppress the formation of coke at the high 
temperatures and long residence times used for toluene dealkylation. 
Higher pressure also reduces the size of the reactor required for a given 
residence time. Higher pressures may also increase the rate of 
dealkylation. 
By reducing the thermal hydrogenator reactor pressure, the synthesis of 
aromatics from paraffin and naphthene hydrocarbons increases. This is the 
reaction sought in the catalytic reforming process. In the reforming 
process the conversion of paraffins and cycloparaffins to aromatics is 
enhanced at reduced pressures. Low pressures also favor increased hydrogen 
production in reformers. A few years ago most reformers operated around 
500 psig to reduce deactivation of the platinum catalyst due to fouling. 
The recent development of more stable catalyst has permitted reformers to 
operate at reduced pressures, i.e. 200 psig or less. 
In the present invention, after maximum conversion has been obtained in a 
catalytic reformer, still further conversion of paraffins to aromatics is 
obtained by charging the reformer effluent to a thermal hydrocracking 
reactor operated at low pressure and low H.sub.2 /oil ratio. Thus while 
few aromatics would be synthesized from paraffins at 600 psig in 
conventional hydrodealkylation, appreciable synthesis can occur at 
selected controlled conditions. 
Another effect of reducing pressure is the production of increased amounts 
of ethane. The rate of thermal hydrocracking of paraffins increases with 
the length of the carbon to carbon chain, with ethane being the most 
thermally stable. At 1,360.degree. F and 25 seconds and 600 psig, most of 
the ethane will react to form methane, and ethane will be less than 25 mol 
% of the total methane and ethane produced. By reducing the temperature to 
1,300.degree. F, the pressure to 200 psig or less, and reducing the 
residence time, ethane and methane can be produced in almost equal molar 
percents. Reducing the H.sub.2 /oil ratio further substantially increases 
the percentage of ethane produced. 
Operations of a thermal hydrocracking reactor at 200 psig pressure and 
1,300.degree. F and a short residence time will give a low rate of toluene 
dealkylation to benzene. Dealkylation of higher aromatics proceeds much 
more rapidly than toluene. The dealkylation rates of various aromatics 
relative to toluene are set forth below. (The data in parenthesis is from 
Betts, Popper, Silsby -- Journal of Applied Chemistry, 7, 497 (1957)). 
______________________________________ 
Aromatics Relative Ratio 
______________________________________ 
Toluene 1 
Ethylbenzene 2.2 
Paraxylene 2.4 (2.9) 
Metaxylene 2.6 (3.5) 
Orthoxylene 4.6 (6.3) 
C.sub.8 Aromatics as group 
2.7 
C.sub.9 Aromatics as group 
4.4 
C.sub.10 Aromatics as group 
10.0 
C.sub.11 Aromatics as group 
17.6 
______________________________________ 
When C.sub.8 or heavier alkyl benzenes are present in feeds there is a very 
large increase in the rate of toluene dealkylation over cases in which 
toluene alone is present. All of the dealkylation reaction rates are 
apparently increased so less severe conditions may be employed with 
mixtures than with pure toluene. The higher reaction rates which occur 
when hydrodealkylating mixtures of aromatics over those experienced in 
dealkylating toluene alone permits the use of much smaller, less expensive 
reactors than would be required if the rate accelerating effect of 
mixtures were not considered in the reactor design. 
It is the purpose of the present process to operate at those conditions 
which will dealkylate C.sub.9 and heavier aromatics to C.sub.8 and 
lighter. At these conditions some C.sub.8 aromatics will also dealkylate, 
but a higher percentage of C.sub.9 aromatics will be dealkylated so the 
percentage yield of mixed xylenes will increase. Since orthoxylene and 
metaxylene dealkylate at a faster rate than paraxylene, the concentration 
of paraxylene in the C.sub.8 fraction will increase. These same conditions 
will also favor the production of additional aromatics from paraffins and 
the production and retention of ethane in lieu of methane to a 
considerable degree. Substantially lower volumes of hydrogen are consumed 
in producing ethane and xylenes versus methane and benzene from a 
catalytic reformer effluent. Hydrogen requirements for the thermal 
hydrogenation reactor are therefore below the volumes produced in the 
catalytic reformer. 
Operation of the reforming unit results in the net production of hydrogen 
which is substantially consumed in the thermal hydrocracking unit under 
the conditions set forth. By controlling the degree of hydrocracking, it 
is possible to ensure that the amount of available hydrogen produced in 
the reformer is sufficient for continuous operation, and thus an external 
source of hydrogen is not required. The product stream most readily 
controlled in accordance with available hydrogen is the butane/heptane 
fraction taken off downstream from the thermal hydrocracking reactor. This 
fraction can be processed closer to extinction by conversion with lighter 
hydrocarbons with adequate hydrogen. Conversely, greater amounts are 
produced as a by-product when sufficient hydrogen is not available. 
Suitable conditions for the thermal hydrocracking reaction includes 
pressures of from about 125 to 275, or up to 375 psig and temperatures 
from about 1200.degree. to 1380.degree. F and residence times from about 5 
to 8 seconds. Recycle hydrogen may be added as discussed below. The 
thermal hydrocracking reactor may be replaced by a catalytic unit which 
operates at a lower temperature but is more difficult to maintain. 
The effluent from the thermal hydrocracking reactor passes through heat 
exchanger 6 and is mixed with rich oil from the absorber 7. After passing 
through coolers 8 the mixture flows to the low pressure flash drum 9. The 
vapors from the drum are compressed in 10, mixed with liquid pumped in 11 
from the drum 9, cooled in 12, and flashed in 13 at a higher pressure. 
About half of the gas from drum 13 is recycled to reformer and the 
remainder of the net product gas is passed through an absorber 7 (and 
optionally a molecular sieve), and into the low temperature hydrogen 
recovery unit 14. The net product gas from the cryogenic separator 14 
contains not over 5% by volume hydrogen and is compressed and may be sold 
as pipe line quality fuel gas or further processed for chemical uses. The 
hydrogen rich vapor is used in desulfurization of the naphtha feedstock 
and as recycle to the catalytic reformer and as quench if needed in the 
THC reactor. Each of the major components separated as overhead from the 
absorber may be recovered separately. Thus as the absorber overhead 
product is cooled, liquid ethane can be separated and as the temperature 
is reduced further liquid methane can be recovered. 
The advantages of this process are illustrated in Table V by comparing the 
composition of the liquid flowing from the high pressure flash drum 13 to 
the stabilizer 16, produced from a 374.degree. F EP naphtha in accordance 
with this invention, with the composition of a reformate from a 
310.degree. F naphtha from the same crude source. The 310.degree. F 
naphtha yields were produced at the severe condition of 125 psig to 
produce the maximum aromatic content for this naphtha versus the less 
severe condition of 200 psig for the 374.degree. F EP. 
TABLE V 
______________________________________ 
Aromatic Content 
LV% of feed 374.degree. F EP 
310.degree. F EP 
Benzene 2.1 5.1 
Toluene 12.7 15.1 
C.sub.8 Aromatics 28.1 18.1 
C.sub.9 Aromatics 12.8 7.4 
C.sub.10 + Aromatics 0.4 0.4 
TOTAL 56.0 46.1 
C.sub.5 + Paraffin & Naphthene 
3.7 32.0 
______________________________________ 
More than 70% of the 3.7% of C.sub.5.sup.+ paraffins and naphthene consists 
of molecules having five and six carbon atoms which will be recovered as 
overhead products in the stabilizer and deisohexanizer. The C.sub.8 
aromatics constitute 50% of the total aromatics produced from the 
374.degree. F EP naphtha but only 39% from the 310.degree. F EP. Also the 
potential for additional C.sub.8 aromatics by disproportionation or 
transalkylation is much higher with the 374.degree. F EP than with the 
310.degree. F EP. Less energy has been expended to this point in the 
process of the invention than in a conventional reforming unit because the 
exothermic heat of thermal hydrocracking has partially offset the 
endothermic heat of reforming. Further the composition of the liquid at 
this point makes downstream processing much easier and less expensive 
(i.e. reduced paraffin, indan, ethyl benzene content). 
Liquid from the high pressure flash drum 13 flows to a stabilizing 
(depentanizing) column 16, a deisohexanizing column 17, and thence to a 
benzene fractionating column 18. The stabilizer separates light 
hydrocarbons overhead which pass to a propane recovery unit 15 from which 
propane and a salable fuel gas are obtained. The stabilizer 16 can be 
eliminated if desired to reduce capital expenditure, in which case propane 
and other light hydrocarbons can be taken overhead from the deisohexanizer 
17. Of course, it is also possible to incorporate additional units or 
eliminate other units in the fractionation chain by incorporating a 
function or step in other more carefully controlled units or by 
eliminating a function or step to produce a less carefully cut product 
fraction. Alternatively, the flow pattern depicted can in effect be 
reversed by taking off the heavier or higher boiling components first and 
sending the remainder downstream for further fractionation. 
A portion of the bottoms product from the deisohexanizer can be recycled to 
the absorber 7 to assist in the separation of fractions in that unit and 
particularly to help break the entrapment of any benzene in the product 
stream to overhead. 
The overhead from the fractionating column 18 is a benzene concentrate 
which, if necessary for very high purity, may be processed in an 
extractive distillation unit 19, such as a Lurgi Distapex, for the 
production of nitration grade benzene and a small quantity of raffinate 
which is combined with the overhead from the deisohexanizer to provide a 
C.sub.4 /C.sub.7 paraffin by-product stream. Table VI sets forth typical 
yield estimates for the extractive distillation unit. 
TABLE VI 
______________________________________ 
EXTRACTIVE DISTILLATION YIELDS 
Charge B/D #/Hr. 
______________________________________ 
C.sub.6 Saturates 1,896 
C.sub.7 Saturates 481 
Benzene from Thermal Hydrocracking (THC) 
5,287 
Benzene from Transalkylation (TA) 
5,693 
Benzene from Isomerization 3,209 
TOTAL 16,566 
Yield 
C.sub.6 /C.sub.7 242 2,377 
Benzene 1,101 14,189 
TOTAL 16,566 
______________________________________ 
Either the deisohexanizer overhead or the raffinate from the benzene 
extractive distillation or both may be recycled to the thermal 
hydrocracking reactor for the production of additional methane and ethane 
to the extent of hydrogen availability. The bottoms from the benzene 
fractionating column are pumped to the toluene column 20. 
The overhead cut from the toluene column is combined with the C.sub.9 
aromatics cut overhead from the rerun column 21 and a portion of the 
bottom cut from the xylene column 22 and charged to the transalkylation 
unit 23. This unit produces primarily xylenes in an equilibrium mixture 
but also some benzene, small amounts of ethyl benzene, and light 
paraffins. 
The transalkylation reaction includes converting C.sub.7 and C.sub.9 (and 
higher) aromatics to C.sub.6 and C.sub.8 (with some C.sub.10 and heavier). 
For example, toluene is passed over a fixed bed of catalyst to induce 
methyl group migration which leads to methylation and demethylation. 
Depending on the type of catalyst used, hydrogen gas may be present to 
depress coke formation. There is little ring destruction or 
hydrodealkylation, and thus minimal hydrogen consumption. Conversion of 
the toluene to benzene and xylenes is about 40% generally, although higher 
conversion is attainable along with increased side reactions. When the 
C.sub.9 and heavier fraction is recycled in equilibrium amounts, benzene 
and xylenes are obtained in virtually equimolar ratio and in close to 
stoichiometric yields. The reaction is an equilibrium system and the 
composition of the product mix depends on the methyl/phenyl ratio in the 
feed, and thus use of increased amounts of C.sub.9 aromatics increases the 
amount of xylenes produced. 
Through the transfer of methyl groups, part of the feed is converted to 
mixed xylenes. Gas, benzene, heavy aromatics and some coke on catatlyst 
are also produced. The catalyst is regenerated by burning off the coke 
deposit. The gas and liquid streams, primarily for recycle, are separated 
in a flash drum which is part of the transalkylation. 
The transalkylation reactor effluent, after heat exchange, flows to the 
stabilizer and fractionating train. Bottoms from the rerun column may be 
charged to the thermal hydrocracking reactor and/or used for fuel. Table 
VII sets forth typical yield estimates for the transalkylation unit. 
TABLE VII 
______________________________________ 
TRANSALKYLATION YIELDS 
Charge Vol. % B/D Mols/Hr. 
Lbs./Hr. 
______________________________________ 
Toluene 48.9 2,539 350.3 32,260 
C.sub.8 Saturates 
0.1 8 0.7 80 
C.sub.9 Aromatics 
49.4 2,566 272.4 32,742 
C.sub.10 Aromatics 
1.5 79 7.4 993 
C.sub.11 Aromatics 
-0.1 6 0.5 74 
TOTAL 100.0 5,198 66,149 
Yield 
Methane 679 
Ethane 500 
Propane 160 
Butane 80 
Benzene 8.5 442 72.9 5,693 
Ethyl Benzene 
3.8 200 23.9 2,541 
Para Xylene 26% 
21.7 1,127 133.9 14,220 
Meta Xylene 50% 
41.7 2,168 258.6 27,468 
Ortho Xylene 24% 
20.0 1,040 126.3 13,413 
Coke 2.02 wt. % 1,395 
TOTAL 66,149 
______________________________________ 
The overhead cut from the xylene column is pumped to the paraxylene 
recovery facility 24 where paraxylene is extracted as a finished product. 
The paraxylene separation is a fractionation process which comprises an 
adsorbent chamber through which the xylene feed and a desorbent (i.e. 
toluene, diethylbenzene) are passed over a fixed bed of a solid adsorbent. 
The desorbent and paraxylene are separated in an extraction column, 
whereas desorbent and raffinate are separated in a raffinate column. The 
adsorbed paraxylene is recovered from the adsorbent by washing it with a 
desorbent liquid having a boiling point different from any of the xylene 
aromatics. The products are separated from the desorbent by fractionation. 
The process arrangement simulates continuous countercurrent flow of 
adsorbent and liquid, without actual movement of the solid. This is 
accomplished with a cycle timer and valves which direct fresh feed and 
stripping liquid to different parts of the bed at predetermined time 
periods. Because of this arrangement, a single bed of adsorbent can be 
used and the flow of feed and products to and from the bed is continuous. 
The other commercially proven method for paraxylene recovery is a 
low-temperature crystallization which may be used in place of the 
adsorbent method. 
Ethyl benzene, metaxylene, and orthoxylene are pumped to the xylene 
isomerization unit 25 where they are partially converted into small 
amounts of gas and benzene and into paraxylene. That stream is recycled to 
the stabilizer and fractionation train. 
In an embodiment of an isomerization procedure a non-equilibrium mixture of 
xylenes and related compounds is driven toward equilibrium -- which under 
the conditions employed may be about 25% paraxylene. 
In this procedure aromatics are hydrogenated over a regenerable catalyst to 
the corresponding naphthene, which is isomerized and then dehydrogenated 
to a different aromatic isomer. The procedure, illustrated generally by 
unit 25, is that the mixed xylenes from the paraxylene recovery unit are 
mixed with hyrogen, preheated and passed through a bed of catalyst. In the 
catalyst bed, the xylenes are isomerized to a near equilibrium mixture of 
orthoxylene, metaxylene, paraxylene, and ethylbenzene. Because of side 
reactions, some compounds both ligher and heavier than xylenes are 
produced also. The reactor outlet is cooled and sent to a flash drum. Part 
of the gas from the flash drum is recycled to the reactor inlet. The rest 
of the gas is purged to fuel. The flash drum liquid is charged to the 
stabilizer where light hydrocarbons are removed. The bottoms are sent to 
the aromatics fractionation section. Table VIII sets forth typical yield 
estimates for the paraxylene extraction and xylene isomerization. 
TABLE VIII 
______________________________________ 
AXYLENE EXTRACTION & 
XYLENE ISOMERIZATION YIELDS 
Charge Mols/Hr. TO- Wt. Lbs. 
From THC TA TAL B/D % /Hr. 
______________________________________ 
Ethyl Benzene 
100.7 23.9 124.6 1,041 13,235 
Paraxylene 
120.1 133.9 254.0 2,138 26,975 
Metaxylene 
280.6 258.6 539.2 4,520 57.268 
Orthoxylene 
170.2 126.3 296.5 2,441 31,488 
C.sub.9 Paraffins 
.2 .2 2 26 
TOTAL 10,142 
100 128,992 
Hydrogen 1,540 
TOTAL 130,532 
Yield 
Methane 15,543 
Ethane 1,800 
Benzene 249 3,209 
Toluene 230 2,917 
Paraxylene 8,485 83 107,063 
TOTAL 130,532 
______________________________________ 
The transalkylation, paraxylene separation and xylene isomerization 
processes are all commercially used and detailed information is available 
from several industry sources. See, for example, From Aromatics To 
Polyester Intermediates by Bergen et al., 1975 UOP Process Division 
Technology Conference, Sept.-Nov. 1975. 
Table IX presents an overall typical yield estimate by weight from the 
extractive distillation, the thermal hydrocracking, the paraxylene 
isomerization and the transalkylation. 
TABLE IX 
______________________________________ 
WEIGHT BALANCE 
Extract PX- Trans 
From: Dist. THC Isom. Alky #/Hr. 
______________________________________ 
H.sub.2 2,898 -1,540 1,358 
C.sub.1 24,872 25,543 679 41,094 
C.sub.2 17,997 1,800 500 20,297 
C.sub.3 13,102 160 13,262 
C.sub.4 10,127 80 10,207 
C.sub.5 4,398 4,398 
C.sub.6 /C.sub.7 
2,377 2,377 
Benzene 14,189 14,189 
Toluene 2,917 2,917 
Paraxylene 107,063 107,063 
Coke 1,395 1,395 
TOTAL 218,557 
Fresh Naphtha Charge 20000 B/D 
218,557 
______________________________________ 
The foregoing process design is adapted to produce a yield of 49% by weight 
of paraxylene (with purity greater than 99.1%) from a heavy naphtha. Based 
on the naphtha in the feedstock, yields of from about 40 to 60% paraxylene 
may be obtained. In addition, nitration grade benzene amounting to 6.5% by 
weight is produced plus 1.3% of toluene for a total yield of 56.8% of 
these aromatics. 
Alternatively, benzene and toluene yields may be increased at the expense 
of paraxylene by altering conditions in the thermal hydrocracking reactor 
particularly by increasing temperature about 50.degree. F and by reducing 
or eliminating flow in the transalkylation unit. The system can be readily 
modified to produce other aromatics. For example orthoxylene can be 
produced from the system disclosed by careful fractionation techniques on 
the xylene stream. Pipeline quality gas production can be increased by 
recycling paraffins to the thermal hydrocracking reactor to the extent 
hydrogen is available. The process of this invention can also be used to 
make ethane as a product from the cryogenic unit by suitable adjustment of 
the condensation and flash conditions. 
In one embodiment of this invention the conventional thermal hydrocracking 
reactor is replaced by the dual reactor system depicted in FIGS. 2 and 3 
having ceramic lined reaction surfaces. The conventional reactors used in 
industry are simple plug flow empty chambers which provide resistance time 
for the reactants and have quench points to control the exothermic 
reaction. In a mixed reactor the kinetic energy required to recycle, 
circulate, or mix the reactor is supplied by dropping the pressure of the 
entire feed through an inlet nozzle. 
In the present invention a somewhat different reactor is designed to 
minimize the pressure drop of the feed and to provide better temperature 
control, better mixing and controlled mixing and some control over the 
hydrogen to feed ratio at various points in the reactor. A dual system 
comprising two reactors, 28 and 29, is used. 
The kinetic energy required for mixing is supplied by the pumped recycle 
liquid 30 (which is heated and injected as a vapor) and the compressed 
hydrogen gas 31 external recycle circuit. The inlet temperature of each of 
these two streams may be precisely controlled. Some of the hot reaction 
products 32 from the bottom of the reactor are recycled back through the 
reactor by means of an eductor 33. The hydrogen/hydrocarbon ratio is lower 
in reactor 28 than in reactor 29 -- promoting the formation of aromatics 
(from paraffins) and ethane in reactor 28 with additional dealkylation 
occurring in reactor 29. The hydrogen recycle feed stream 31 is injected 
into and out of the upper portion of reactor 28 to provide an eductor 
effect which draws up and provides mixing with the fresh feed and recycle 
stream 30. The temperature of the hydrogen recycle stream is used to help 
control the temperature in the reactors. Control is also facilitated by 
the use of two small reactors in place of one large unit. This arrangement 
eliminates the elaborate quench systems which have been used and still 
maintains close control over temperature and promotes the desired 
reactions. The tangential entry flow pattern of the streams to the 
reactors is illustrated in FIG. 3. 
This invention has been described in terms of specific embodiments set 
forth in detail, but it should be understood that these are by way of 
illustration only and that the invention is not necessarily limited 
thereto. Modifications and variations will be apparent from this 
disclosure and may be resorted to without departing from the spirit of 
this invention, as those skilled in the art will readily understand. 
Accordingly, such variations and modifications of the disclosed process 
are considered to be within the purview and scope of this invention and 
the following claims.