U.S. patent application number 13/140458 was filed with the patent office on 2011-12-15 for process for preparing benzene.
This patent application is currently assigned to BAYER TECHNOLOGY SERVICES GMBH. Invention is credited to Evin Hizaler Hoffmann, Bharat Marwaha, Leslaw Mleczko, Ralph Schellen, Stephan Schubert.
Application Number | 20110303581 13/140458 |
Document ID | / |
Family ID | 42167570 |
Filed Date | 2011-12-15 |
United States Patent
Application |
20110303581 |
Kind Code |
A1 |
Schellen; Ralph ; et
al. |
December 15, 2011 |
PROCESS FOR PREPARING BENZENE
Abstract
The present invention relates to a process for the endothermic,
catalytic gas phase reaction of naphtha with hydrogen to form
benzene, in which the reaction is carried out in 5 to 12 serial
reaction zones under adiabatic conditions.
Inventors: |
Schellen; Ralph; (Dormagen,
DE) ; Hizaler Hoffmann; Evin; (Koln, DE) ;
Mleczko; Leslaw; (Dormagen, DE) ; Schubert;
Stephan; (League City, TX) ; Marwaha; Bharat;
(Pearland, TX) |
Assignee: |
BAYER TECHNOLOGY SERVICES
GMBH
Leverkusen
DE
|
Family ID: |
42167570 |
Appl. No.: |
13/140458 |
Filed: |
December 4, 2009 |
PCT Filed: |
December 4, 2009 |
PCT NO: |
PCT/EP09/08672 |
371 Date: |
September 2, 2011 |
Current U.S.
Class: |
208/49 |
Current CPC
Class: |
C07C 2523/36 20130101;
C07C 5/367 20130101; C07C 2521/04 20130101; C10G 35/04 20130101;
C07C 5/417 20130101; C07C 2523/42 20130101; C07C 5/417 20130101;
C07C 15/04 20130101; C07C 15/04 20130101; C10G 2400/30 20130101;
C07C 5/367 20130101 |
Class at
Publication: |
208/49 |
International
Class: |
C10G 65/02 20060101
C10G065/02 |
Foreign Application Data
Date |
Code |
Application Number |
Dec 20, 2008 |
DE |
10 2008 064 276.2 |
Claims
1. Process for preparing benzene from naphtha in the presence of
hydrogen in an endothermic, heterogeneously catalytic gas phase
reaction, which comprises 5 to 12 serial reaction zones with
adiabatic conditions.
2. Process according to claim 1, wherein the conversion takes place
in 6 to 10 serial reaction zones.
3. Process according to claim 1, wherein the entry temperature of
the process gas entering the first reaction zone is 740 to 790
K.
4. Process according to claim 1, wherein the absolute pressure at
the entry of the first reaction zone is between 10 and 40 bar.
5. Process according to claim 1, wherein the residence time of the
process gas in all reaction zones is between 0.5 and 30 s.
6. Process according to claim 1, wherein the catalysts are present
in a fixed bed arrangement.
7. Process according to claim 6, wherein the catalysts are present
in the form of monoliths.
8. Process according to claim 7, wherein the monolith comprises
channels having a diameter of 0.1 to 3 mm.
9. Process according to claim 1, wherein the catalysts are present
in beds of particles having average particles sizes of 1 to 10
mm.
10. Process according to claim 1, wherein at least one reaction
zone is followed by at least one heat exchange zone through which
the process gas is passed.
11. Process according to claim 10, wherein each reaction zone is
followed by at least one heat exchange zone through which the
process gas is passed.
12. Process according to claim 1, wherein between a reaction zone
and a heat exchange zone there is at least one heat insulation
zone.
13. Process according to claim 12, wherein around each reaction
zone there is a heat insulation zone.
Description
[0001] The present invention relates to a process for the
endothermic, catalytic gas phase reaction of naphtha with hydrogen
to form benzene, in which the reaction is carried out in 5 to 12
serial reaction zones under adiabatic conditions.
[0002] Naphtha is an untreated petroleum distillate from the
refining of petroleum or natural gas, and one of the products
typically recovered from it is benzene. Benzene in turn is a key
starting material for many further petrochemicals.
[0003] For instance, benzene is used in the Chemical Industry for
the synthesis of numerous compounds, such as, for example, aniline,
styrene, nylon, synthetic rubber, plastics, detergents,
insecticides, dyes and numerous further substances. Also obtained,
by substitution, are numerous aromatics, such as, for example,
phenol, nitrobenzene, aniline, chlorobenzene, hydroquinone and
picric acid.
[0004] A further product, though one which has now taken a back
seat for environmental reasons, is the use of the benzene as a fuel
for internal combustion engines operating in accordance with the
Otto cycle.
[0005] The key reactions in the preparation of benzene from naphtha
are set out in the formulae below (I to IV). Formula (I) relates to
the conversion of cyclohexane as a fraction of the naphtha, to form
benzene, and is very endothermic, and formula (II) relates to the
conversion of hexane to form cyclohexane, which can be reformed in
turn into benzene in accordance with formula (I). The secondary
reactions of formulae (III) and (IV), which may likewise take place
during the preparation of benzene, are shown as well, and are
especially exothermic reactions.
C.sub.6H.sub.12C.sub.6H.sub.6+3.H.sub.2 (I)
C.sub.6H.sub.14C.sub.6H.sub.12+H.sub.2 (II)
C.sub.6H.sub.12+2.H.sub.2.fwdarw.2.C.sub.3H.sub.8 (III)
C.sub.6H.sub.14+H.sub.2.fwdarw.2.C.sub.3H.sub.8 (IV)
[0006] The reaction according to formula (I), like that according
to formula (II), is equilibrium-limited.
[0007] The benzene obtained from the reaction according to formula
(I) forms a key starting product for further reaction to give, for
example, the abovementioned products.
[0008] The controlled supply of heat in processes for obtaining
benzene is important, since the position of the equilibrium in the
aforementioned reaction according to formula (I) is heavily
dependent on the temperature of the reaction zone, and it is
therefore possible to control the yields and/or selectivities with
respect to benzene by this means. In particular it is therefore
possible at least partly to suppress the unwanted secondary
reactions of formulae (III-IV).
[0009] An uncontrolled drop in temperature as a result of the
endothermic reaction according to formula (I) may therefore promote
the formation of more or less large quantities of cyclohexane
(according to formula I by back-reaction) and/or propane (according
to formula III) and/or hexane (according to formula II by
back-reaction), which is a disadvantage for the subsequent use of
the benzene, since such secondary constituents must first be
separated off.
[0010] It is therefore advantageous to keep the temperature of the
reaction zones in the course of the process controlled at a level
which allows rapid conversion with a minimization of the secondary
reactions.
[0011] Accordingly, EP 0 601 398 A1 discloses a key influence on
yield and conversion to target product exerted in the preparation
of BTX aromatics (benzene, toluene and xylene) by the temperature
level and by the catalyst employed. In accordance with the
disclosure content of EP 0 601 398 A1, the temperature level at
which the reaction is to be performed is essentially determined by
the nature and composition of the naphtha used, which is typically
characterized by its boiling point. This underlines the importance
of precise temperature control in such processes.
[0012] EP 0 601 398 A1 also discloses how it is now customary to
perform catalytic reformation processes in a plurality of reactors,
connected in series, containing catalysts in the form of a fixed
bed.
[0013] EP 0 601 398 A1 discloses an isothermal procedure using a
salt bath by means of which the temperatures of approximately
500.degree. C. that are disclosed in the process are brought about
in the reaction zone. An adiabatic regime is not disclosed. The
catalyst used in the process disclosed in EP 0 601 398 A1 is
composed of a support material, which is preferably alumina, on
which there is located a layer of a platinum group metal with a
promoter metal from group WB of the Periodic Table of the
Elements.
[0014] The process as disclosed in EP 0 601 398 A1 is
disadvantageous on account of the fact that the isothermal
procedure disclosed is extremely complicated and therefore very
expensive. In the production of industrial chemicals in particular,
which includes the production of benzene, however, even slight
process disadvantages have severe consequences for the economics of
the process as a whole, as is also disclosed by EP 0 601 398
A1.
[0015] The possibility of an adiabatic regime is disclosed by J.
Ancheyta-Juarez et al. in "Modeling and simulation of four
catalytic reactors in series for naphtha reforming" in Energy &
Fuels (2001) 15: 887-893.
[0016] Thus J. Ancheyta-Juarez et al. disclose how it can be
advantageous to perform the reaction of naphtha to form (among
other products) benzene in three to four, especially four, serial
reaction zones, with the possibility of intermediate cooling
between the aforementioned reaction zones.
[0017] The yields of benzene (A.sub.6) which can be achieved by
means of the process presented in the disclosure by J.
Ancheyta-Juarez et al. are very low, with a fraction of only about
4 mol % as a proportion of the reaction product, thus making the
process disadvantageous.
[0018] EP 1 251 951 (B1) discloses an apparatus and the possibility
of conducting chemical reactions in the apparatus, the apparatus
being characterized by a cascade of mutually contacting reaction
zones and heat exchanger devices which are integrated materially
with one another. The process to be conducted therein is
characterized, therefore, by the contact of the various reaction
zones with a respective heat exchanger device, in the form of a
cascade. A disclosure as to the possibility of using the device and
the process for preparing benzene is absent.
[0019] It remains unclear, then, as to how, on the basis of the
disclosure content of EP 1 251 951 (B1), a reaction of this kind is
to be carried out by means of the apparatus and of the process
performed therein. In particular there is no disclosure of a
process comprising endothermic reactions.
[0020] Furthermore, for reasons of consistency, it must be assumed
that the process disclosed in EP 1 251 951 (B1) is performed in an
apparatus which is identical or similar to the disclosure relating
to the apparatus. The result of this is that, as a consequence of
the extensive contact between the heat exchange zones and the
reaction zones, as per the disclosure, a significant amount of heat
takes place as a result of thermal conduction between the reaction
zones and the adjacent heat exchange zones.
[0021] The disclosure concerning the oscillating temperature
profile can only be understood, then, to mean that the temperature
peaks found here would be sharper in the absence of this contact. A
further indicator of this is the exponential increase in the
disclosed temperature profiles between the individual temperature
peaks. These indicate that there is a certain heat sink present in
each reaction zone, with a marked but limited capacity, which is
able to reduce the temperature increase in said zone. It is never
possible to rule out a certain dissipation of heat (by radiation,
for example); however, in the case of a reduction in the possible
dissipation of heat from the reaction zone, a linear or degressive
temperature profile would be suggested, since there is no
subsequent metering of reactants and hence, after their consumption
by exothermic reaction, the reaction would become increasingly slow
and hence the heat produced would go down.
[0022] EP 1 251 951 (B1) thus discloses multi-stage processes in
cascades of reaction zones from which heat is taken off in an
undefined quantity by means of thermal conduction. The process
disclosed, therefore, is not adiabatic and is disadvantageous
insofar as precise temperature control of the reaction is
impossible. This applies especially to the undisclosed possibility
of an endothermic reaction in the reaction zones.
[0023] On the basis of the prior art, therefore, it would be
advantageous to provide a process for preparing benzene that can be
carried out in simple reaction apparatus and that enables precise,
simple temperature control of the endothermic process, thereby
allowing high conversions in conjunction with very high purities of
the product, while meeting desired yields and/or selectivities.
Said simple reaction apparatus would be readily transposable to the
industrial scale, and inexpensive and robust in all sizes.
[0024] As has just been shown, there have to date been no suitable
processes apparent that allow this for the endothermic, catalytic
gas phase reaction of naphtha to form benzene.
[0025] The object is therefore that of providing a process for
endothermic, catalytic gas phase reaction of naphtha to benzene
which can be carried out with precise temperature control in simple
reaction apparatus and which as a result allows high conversions in
conjunction with high product purities.
[0026] It has surprisingly been found that a process for preparing
benzene from naphtha in the presence of hydrogen in an endothermic,
heterogeneously catalytic gas phase reaction, characterized in that
it comprises 5 to 12 serial reaction zones with adiabatic
conditions, is able to achieve this object.
[0027] In connection with the present invention, benzene is a
process gas substantially comprising benzene. The benzene may also
comprise fractions of hydrogen and further hydrocarbons.
[0028] In connection with the present invention, further
hydrocarbons are compounds present in the form of process gas
composed of carbon, hydrogen and possibly oxygen. Essentially,
however, such hydrocarbons are composed of carbon and hydrogen.
Such hydrocarbons are typically either those which are introduced
into the process of the invention as further constituents of the
naphtha, or those which are formed as a result of secondary
reactions in the course of the process of the invention, as for
instance by the reactions according to formulae (III and IV).
[0029] Non-exhaustive examples of hydrocarbons which are introduced
into the process of the invention as further constituents of the
naphtha are naphthalene, isopentane and toluene, for instance.
[0030] Non-exhaustive examples of hydrocarbons which are formed in
the course of the process of the invention by secondary reactions,
as for instance by the reactions according to formulae (III and
IV), are hexane, cyclohexane and propane, for instance.
[0031] Naphtha identifies a mixture of hydrocarbons in the form of
a process gas, as is general knowledge to the person skilled in the
art. In connection with the process of the invention, naphtha is
preferably a mixture of hydrocarbons substantially comprising
cyclohexane.
[0032] In connection with the present invention, hydrogen is a
process gas which substantially comprises hydrogen. This hydrogen
may be formed, for instance, by the reactions according to formulae
(I and II), or else may be supplied as process gas to the
process.
[0033] The supplying of hydrogen as a process gas into the process
of the invention is preferred. With particular preference,
preheated hydrogen is supplied as process gas to the process.
[0034] Such supplying of hydrogen in particular is advantageous in
that it allows the hydrogen to be used as a heat transfer medium in
the process, for controlling the temperature. Furthermore, the
hydrogen prevents deposits of carbon products on the catalyst
surfaces of the catalysts located in the reaction zones
(coking).
[0035] The identification "substantially" refers, in connection
with the present invention, to a mass fraction and/or a molar
fraction of at least 80%.
[0036] The naphtha used in the process of the invention, its
constituents, the hydrogen, the benzene and also the products of
the process of the invention are also referred to below
collectively as process gases.
[0037] It follows from this that the entire process of the
invention is performed in the gas phase. If substances used in the
process, such as the hydrocarbons, for instance, are not in gaseous
form at room temperature (23.degree. C.) and ambient pressure (1013
hPa), it can be assumed below that, before or during their use in
the process of the invention, such substances will be converted
into the gas phase by an increase in temperature and/or reduction
in pressure.
[0038] Besides the substantial components of the process gases,
they may also comprise secondary components. Non-exhaustive
examples of secondary components which may be present in the
process gases are argon, nitrogen and/or carbon dioxide, for
instance.
[0039] In accordance with the invention the implementation of the
process under adiabatic conditions means that substantially neither
heat is actively supplied nor heat withdrawn from the outside
to/from the reaction zone. It is common knowledge that complete
insulation from ingress or egress of heat is possible only by
complete evacuation, with the possibility of heat transfer by
radiation being ruled out. In connection with the present
invention, therefore, adiabatic means that no measures are taken to
supply or remove heat.
[0040] In one alternative embodiment of the process of the
invention, however, heat transfer may be reduced, for example, by
insulation using conventional insulating means, such as polystyrene
insulants, for example, or else by sufficiently large distances
from heat sinks or heat sources, the insulation means being
air.
[0041] An advantage of the adiabatic regime of 5 to 12 serial
reaction zones in accordance with the invention as compared with a
non-adiabatic regime is that in the reaction zones there is no need
to provide means for heat removal, a fact which results in a
considerable simplification in construction. As a result, in
particular, there are simplifications in the manufacturing of the
reactor and also in the scaleability of the process, and there is
an increase in the reaction conversions.
[0042] A further advantage of the process of the invention is the
possibility of very precise temperature control, as a result of the
narrow staggering of adiabatic reaction zones. By this means it is
possible in each reaction zone to set and control a temperature
which is advantageous in the progress of the reaction.
[0043] The catalysts used in the process of the invention are
typically catalysts composed of a material which as well as its
catalytic activity for the reaction according to formula (I) is
characterized by sufficient chemical resistance under the
conditions of the process and also by a high specific surface
area.
[0044] Catalyst materials which are characterized by such chemical
resistance under the conditions of the process are, for example,
catalysts comprising platinum and/or rhenium.
[0045] Preferred catalyst materials are composed of equal weight
fractions of rhenium and platinum.
[0046] These catalysts may be applied on support materials. Such
support materials typically comprise alumina and/or titanium
dioxide. Preference is given to alumina support materials.
[0047] Particularly preferred catalysts are composed of rhenium and
platinum applied at the same weight fraction on an alumina support.
Methods of producing such catalysts are general knowledge to a
person skilled in the art, from EP 0 601 398 A1, for instance.
[0048] Specific surface area in connection with the present
invention identifies the surface area of the catalyst material
which can be reached by the process gas, based on the mass of
catalyst material employed.
[0049] A high specific surface area is a specific surface area of
at least 1 m.sup.2/g, preferably of at least 10 m.sup.2/g.
[0050] The catalysts of the invention are located in the reaction
zones in each case and may be present in all conventional
presentation forms, e.g. fixed bed, moving bed.
[0051] The presentation form is preferably that of a fixed bed.
[0052] The fixed bed arrangement comprises a catalyst bed in the
actual sense, i.e. loose, supported or unsupported catalyst in any
desired form, and also in the form of suitable packings. The term
catalyst bed as used herein also encompasses coherent regions of
suitable packings on a support material or structured catalyst
support. Examples of such would include ceramic honeycomb supports
for coating, having comparatively high geometric surface areas, or
corrugated layers of metal wire mesh with catalyst granules, for
example, immobilized thereon. In connection with the present
invention, the presence of the catalyst in monolithic form is
viewed as a special form of packing.
[0053] Where a fixed bed arrangement of the catalyst is used, the
catalyst is preferably in beds of particles having average
particles sizes of 1 to 10 mm, preferably 2 to 8 mm, more
preferably of 3 to 7 mm.
[0054] Likewise with preference the catalyst in the case of a fixed
bed arrangement is in monolithic form. In the case of a fixed bed
arrangement particular preference is given to a monolithic catalyst
which comprises the aforementioned metals, rhenium and platinum, in
equal weight fractions on an alumina support.
[0055] Likewise particularly preferred is a fixed bed arrangement
having particle beds, having average particles sizes of 1 to 10 mm,
preferably 2 to 8 mm, more preferably of 3 to 7 mm, the particles
being alumina particles to which the aforementioned metals, rhenium
and platinum, have been applied in equal weight fractions.
[0056] If a catalyst is used in monolithic form in the reaction
zones, then, in a preferred development of the invention, the
catalyst present in monolithic form is provided with channels
through which the process gases flow. Typically the channels have a
diameter of 0.1 to 3 mm, preferably a diameter of 0.2 to 2 mm, more
preferably of 0.5 to 1.5 mm.
[0057] Where a moving bed arrangement of the catalyst is used, the
catalyst preferably takes the form of loose beds of particles, of
the kind already described in connection with the fixed bed
arrangement.
[0058] Beds of such particles are advantageous because the size of
the particles have a high specific surface area of the catalyst
material in relation to the process gases and it is therefore
possible to achieve a high conversion rate. Accordingly, the
limitation of mass transport in the reaction by diffusion can be
minimized. At the same time, however, the particles are also not so
small that increased pressure drops occur disproportionately when
the gases flow through the fixed bed. The ranges of particle sizes
specified in the preferred embodiment of the process, comprising a
reaction in a fixed bed, are therefore an optimum between the
achievable conversion from the reactions according to formulae (I
and II) and the pressure drop generated when the process is
implemented. Pressure drop is coupled directly with the necessary
energy, in the form of compressor output, and so a
superproportional increase in the latter would result in an
uneconomic process regime.
[0059] In one preferred embodiment of the process of the invention
the conversion takes place in 6 to 10, more preferably 6 to 8,
serial reaction zones.
[0060] A preferred further embodiment of the process is
characterized in that the process gas emerging from at least one
reaction zone is subsequently passed through at least one heat
exchange zone downstream of said reaction zone.
[0061] In one particularly preferred further embodiment of the
process each reaction zone is followed by at least one, preferably
exactly one, heat exchange zone through which the process gas
emerging from the reaction zone is passed.
[0062] These reaction zones may be disposed either in one reactor
or, in divided form, in two or more reactors. The arrangement of
the reaction zones in one reactor leads to a reduction in the
number of apparatuses used.
[0063] The individual reaction zones and heat exchange zones may
also be arranged in one reactor or, in divided form, in any desired
combinations of reaction zones with heat exchange zones in two or
more reactors.
[0064] Where reaction zones and heat exchange zones are present in
one reactor, then, in an alternative embodiment of the invention,
there is a heat insulation zone located between these zones in
order to allow the adiabatic operation of the reaction zone to be
maintained.
[0065] In addition it is possible for certain of the serial
reaction zones to be replaced or supplemented, independently of one
another, by one or more parallel reaction zones. The use of
parallel reaction zones makes it possible in particular to exchange
or supplement the zones during ongoing, continuous operation of the
process overall.
[0066] Parallel and serial reaction zones may in particular also be
combined with one another. With particular reference, however, the
process of the invention features exclusively serial reaction
zones.
[0067] The reactors used preferably in the process of the invention
may be composed of simple vessels with one or more reaction zones,
of the kind described, for example, in Ullmann's Encyclopedia of
Industrial Chemistry (Fifth, Completely Revised Edition, Vol. B4,
pages 95-104, pages 210-216), it being possible for heat insulation
zones to be provided additionally between each of the individual
reaction zones and/or heat exchange zones.
[0068] In one alternative embodiment of the process, therefore,
there is at least one heat insulation zone located between a
reaction zone and a heat exchange zone. Preferably there is a heat
insulation zone located around each reaction zone.
[0069] The catalysts or fixed beds thereof are mounted in a
conventional way on or between gas-permeable walls encompassing the
reaction zone of the reactor. In the case of thin fixed beds in
particular, technical devices for uniform gas distribution may be
fitted upstream of the catalyst beds. These devices may be
perforated plates, bubble trays, valve trays or other internals
which, by generating a low but uniform pressure drop, produce
uniform entry of the process gas into the fixed bed.
[0070] In one preferred embodiment of the process the entry
temperature of the process gas entering one reaction zone is from
740 to 790 K, preferably from 750 to 780 K, more preferably from
755 to 775 K.
[0071] In a further preferred embodiment of the process the
absolute pressure at the entry of the first reaction zone is
between 10 and 40 bar, preferably between 15 and 35 bar, more
preferably between 20 and 30 bar.
[0072] In yet another preferred embodiment of the process the
residence time of the process gas in all the reaction zones
together is between 0.5 and 30 s, preferably between 1 and 20 s,
more preferably between 5 and 15 s.
[0073] The naphtha and, where appropriate, the hydrogen are
preferably supplied only ahead of the first reaction zone. This has
the advantage that the whole of the process gas is available for
the accommodation of heat of reaction in all the reaction zones.
Such a procedure additionally enables the space-time yield to be
increased, or the mass of catalyst required to be reduced. It is,
however, also possible to meter naphtha and, where appropriate,
hydrogen into the process gas ahead of one or more of the reaction
zones that follow the first reaction zone, if needed. The supply of
these process gases between the reaction zone is an additional way
of controlling the temperature of the conversion, if they are
preheated.
[0074] In preferred embodiments of the process of the invention the
molar ratio of hydrogen to hydrocarbons present in the naphtha is
set in ranges from 3 to 9, preferably from 4 to 8, more preferably
from 5 to 7 mol of hydrogen per mole of hydrocarbon in the
naphtha.
[0075] The advantages of such supply of hydrogen have already been
elucidated. They apply in particular in connection with the supply
of an excess.
[0076] The person skilled in the art is aware of suitable means for
determining the molar amounts of hydrocarbons in a process gas,
such as naphtha. One non-exhaustive example is quantitative
analysis by means of gas chromatography. If the molar composition
of the naphtha process gas is known, the molar ratio of hydrogen to
it can be set by simple setting of the volume flow ratio of the
naphtha and hydrogen process gases.
[0077] In a further preferred embodiment of the process of the
invention the process gas is heated after at least one of the
reaction zones used, more preferably after each reaction zone. This
is done by passing the process gas, following exit from a reaction
zone, through one or more of the abovementioned heat exchange zones
which are located downstream of the respective reaction zones.
These zones may be configured as heat exchange zones in the form of
heat exchangers known to the person skilled in the art, such as
shell-and-tube, plate, annular-groove, spiral, ribbed-tube or
micro-type heat exchangers, for example. The heat exchangers are
preferably microstructured heat exchangers.
[0078] Microstructured in connection with the present invention
means that the heat exchanger, for the purpose of heat transfer,
comprises fluid-carrying channels which are characterized in that
they have a hydraulic diameter of between 50 .mu.m and 5 mm. The
hydraulic diameter is calculated from four times the flow-traversed
cross-sectional area of the fluid-carrying channel, divided by the
circumference of the channel.
[0079] In one particular embodiment of the process the process gas
is heated in the heat exchange zones by condensation of a heat
transfer medium.
[0080] Within this particular embodiment it is preferred to perform
condensation, preferably partial condensation, on the side of the
heating medium in the heat exchangers which constitute the heat
exchange zones.
[0081] Partial condensation in connection with the present
invention means a condensation in which the heating medium used is
a substance in the form of a gas/liquid mixture, and in which,
following heat transfer in the heat exchanger, this substance is
still in the form of a gas/liquid mixture.
[0082] Performing a condensation is particularly advantageous since
it means that the coefficient of heat transfer to the process gases
from the heating medium that can be achieved becomes particularly
high and hence that it is possible to achieve efficient
heating.
[0083] Performing a partial condensation is particularly
advantageous because it means that the delivery of heat by the
heating medium no longer results in a temperature change to the
heating medium, but instead merely shifts the gas/liquid
equilibrium. As a result of this, the process gas is heated towards
a constant temperature over the entire heat exchange zone. This in
turn reliably prevents the incidence of radial temperature profiles
in the flow of the process gases, thereby improving the control of
the reaction temperatures in the reaction zones and, in particular,
preventing the development of instances of local overheating as a
result of radial temperature profiles.
[0084] In an alternative embodiment, instead of a
condensation/partial condensation, it is also possible to provide a
mixing zone before the entry of a reaction zone, in order to unify
any radial temperature profiles formed in the course of heating in
the flow of the process gases by mixing transverse to the main flow
direction.
[0085] In one preferred embodiment of the process the succession of
reaction zones are operated with an average temperature rising or
falling from reaction zone to reaction zone. This means that,
within a sequence of reaction zones, the temperature may be made
both to rise and to fall from reaction zone to reaction zone. This
can be brought about, for example, by means of control of the heat
exchange zones inserted between the reaction zone. Further
possibilities for setting the average temperature are described
below.
[0086] The thickness of the flow-traversed reaction zones may be
the same or different and is a function of laws which are general
knowledge to the person skilled in the art and relate to the
above-described residence time and the volumes of process gas
processed in each case. The mass flows of process gas that can be
processed in accordance with the invention, relative to the mass of
catalyst used (and also called WHSV, Weight-Hourly Space Velocity),
is typically between 28 and 42 h.sup.-1, preferably between 30 and
40 h.sup.-1, more preferably between 33 and 38 h.sup.-1.
[0087] The maximum exit temperature of the process gas from the
first reaction zone is typically in the region of the entry
temperature, since the reactions according to formulae (III) and
(W) are exothermic reactions. Particularly in the case of exit from
the final reactions, in which a large amount of benzene has already
been formed and therefore the especially endothermic reaction
according to formula (I) is losing its influence, these
temperatures may also be situated within a range from 770 to 820 K,
preferably from 775 to 795 K, more preferably from 780 to 785
K.
[0088] The person skilled in the art is able freely to determine
the entry temperature of the subsequent reaction zones by means of
the measures below in accordance with the process of the
invention.
[0089] Temperature control in the reaction zones is accomplished
preferably by at least one of the following measures: sizing of the
adiabatic reaction zone, control of the heat supply between the
reaction zones, addition of further process gas between the
reaction zones, molar ratio of reactants/excess of hydrogen used,
addition of secondary constituents, especially nitrogen, carbon
dioxide, ahead of and/or between the reaction zones.
[0090] The composition of the catalysts in the reaction zones of
the invention may be the same or different. In one preferred
embodiment the same catalysts are used in each reaction zone. An
advantageous alternative is to use different catalysts in the
individual reaction zones.
[0091] Thus it is possible in particular in the first reaction
zone, when the concentration of the reactants is still high, to use
a less active catalyst, and to increase the activity of the
catalyst from reaction zone to reaction zone in the further
reaction zones. The catalyst activity can also be controlled by
dilution with inert materials and/or support material.
[0092] With the process of the invention it is possible to prepare
1 kg/h to 50 kg/h, preferably 5 kg/h to 30 kg/h, more preferably 10
kg/h to 20 kg/h of benzene per kg of catalyst.
[0093] The process of the invention is therefore distinguished by
high space-time yields, in conjunction with a reduction in
apparatus sizes and also with a simplification of the apparatus
and/or reactors. This surprisingly high space-time yield is made
possible through the interaction of the inventive and preferred
embodiments of the new process. The interaction of staggered
adiabatic reaction zones with interposed heat exchange zones and
the defined residence times, in particular, allows precise control
of the process and the resulting high space-time yields, and also a
reduction in the by-products formed, such as carbon dioxide, for
instance.
[0094] The present invention is illustrated with reference to the
figures, though without being restricted thereto.
[0095] FIG. 1 shows reactor temperature (T) and molar mass flows of
benzene (U) over a length (L) of 11 m of reaction zones each with
downstream heat exchange zones (in accordance with Example 1), the
lengths of the heat exchange zones being assumed ideally to be
zero, since no conversion is to take place here.
[0096] The present invention is further illustrated by the
following example, but without being limited thereto.
EXAMPLES
[0097] Gaseous naphtha and hydrogen are supplied to the process as
process gases in a molar ratio of 7.77. The process is operated in
a total of six fixed catalyst beds of rhenium and platinum, each at
0.29% by weight, on an alumina support, in other words in six
reaction zones.
[0098] After each reaction zone there is a heat exchange zone in
which the exiting process gas is heated again before entering the
next reaction zone.
[0099] The absolute entry pressure of the process gas directly
ahead of the first reaction zone is 25 bar. The length of the fixed
catalyst beds, and therefore of the reaction zones, varies from
reaction zone to reaction zone, beginning from 0.15 m in the first
reaction zone through to 6 m in the sixth reaction zone. The
precise links of the reaction zone are summarized in Table 1. The
activity of the catalyst used is unvarying over the reaction zones.
No process gas is metered in ahead of the individual reaction
zones. The WHSV is 35 h.sup.-1.
TABLE-US-00001 TABLE 1 Lengths of the reaction zones Reaction
Length zone [#] [m] 1 0.15 2 0.35 3 1 4 1.5 5 2 6 6 .SIGMA.
11.0
[0100] The results are shown in FIG. 1. Varying the cumulative
length of the reaction zones is plotted on the x-axis, so that it
is possible to see a spatial course of the developments in the
process; the heat exchange zones are disregarded. On the left-hand,
y-axis, the temperature of the process gas is indicated. The
temperature profile over the individual reaction zones is depicted
as a thick, continuous line. As a result of the idealized
assumption of the length of the heat exchange zones as 0 m, there
are discontinuities in the temperature profile. On the right-hand
y-axis the cumulative molar flow of benzene in the process gas over
the reaction path is indicated. Its profile over said path is
depicted as a thin continuous line.
[0101] It can be seen that the entry temperature of the process gas
ahead of the first reaction zone is approximately 775 K. As a
result of the substantially endothermic reaction to form benzene
under adiabatic conditions the temperature in the first reaction
zone drops to about 760 K, before in the downstream heat exchange
zone the process gas is reheated to the aforementioned 775 K. As a
result of endothermic adiabatic reaction, the temperature in the
second reaction zone drops to about 750 K. The sequence of cooling
as a result of endothermic, adiabatic reaction, and heating
continues, with changes in exit temperatures, after the respective
reaction zones, the entry temperature being re-established at the
desired 775 K in each of the heat exchange zones.
[0102] A conversion of cyclohexane and hexane of approximately 60%
is obtained. The space-time yield achieved, based on the mass of
catalyst employed, is approximately 15
kg.sub.benzene/kg.sub.cath.
* * * * *