Hydrogenation process

In a process for the hydrogenation of multiply unsaturated C.sub.2 -C.sub.8 hydrocarbons, in particular C.sub.2 -C.sub.8 -alkynes and/or C.sub.4 -C.sub.8 -alkynenes and/or C.sub.4 -C.sub.8 -alkadienes in fluids comprising these by contact with a catalyst packing in the presence of free hydrogen, the catalyst packing can be produced by applying at least one substance which is active as catalyst and/or promoter to woven or knitted meshes or foils as support material.

The present invention relates to a process for hydrogenating C.sub.2 
-C.sub.8 -alkynes and/or C.sub.4 -C.sub.8 -alkynenes and/or C.sub.4 
-C.sub.8 -alkadienes in fluids comprising these. Alkynes, eg. acetylene, 
and dienes are undesired materials in many industrial syntheses because of 
their tendency to polymerize and their pronounced tendency to form 
complexes with transition metals. They sometimes have a very strong 
adverse effect on the catalysts used in these reactions. Thus, for 
example, the acetylene present in the C.sub.2 stream of a steam cracker 
interferes with the polymerization of ethylene, so that the acetylene 
content of the C.sub.2 stream has to be kept very small, preferably less 
than 1 ppm. The C.sub.3 stream of a steam cracker, which comprises not 
only propylene but also from 2 to 3% of propadiene (PD) and about the same 
amount of propyne (methylacetylene, MA), also has to be purified before 
polymerization to give polypropylene. The typical content of multiply 
unsaturated hydrocarbons is here from about 4 to 6% by weight. A reduction 
of this content to a maximum of 10 ppm should preferably be achieved. The 
C.sub.4 stream of a steam cracker also contains up to 70% of multiply 
unsaturated hydrocarbons. These are mainly butadiene, vinylacetylene and 
ethylacetylene. The total content of multiply unsaturated hydrocarbons 
should be reduced to less than 20 ppm, preferably a maximum of 10 ppm. 
This is achieved in industry by selective hydrogenation of the hydrocarbon 
streams over heterogeneous noble metal catalysts on ceramic supports. High 
demands are here placed on the hydrogenation catalysts used in respect of 
their selectivity and activity, since a very complete hydrogenation of the 
multiply unsaturated hydrocarbons without loss of monounsaturated 
hydrocarbons such as ethylene, propene or butenes is to be achieved. 
In some cases, a raw C.sub.4 stream from a steam cracker, which contains 
from about 40 to 60% by weight of butadiene, is to be selectively 
hydrogenated to form butenes in as high as possible a yield. In this case 
too, industrial heterogeneous noble metal catalysts on ceramic supports 
are used. For these applications, use is made of promoted or unpromoted 
noble metal catalysts on ceramic supports usually with palladium as active 
component in an amount of from 0.01 to 1% by weight. 
In the known processes, carbon monoxide is often mixed into the reaction 
mixture for the hydrogenation of acetylene in order to increase the 
selectivity of the catalyst. The disadvantage of this method is that the 
selectivity-increasing action of the carbon monoxide is strongly 
temperature-dependent. Large temperature gradients in the catalyst bed 
therefore result in a worsening of the selectivity. In addition, the 
relatively high working temperatures which are necessary when carbon 
monoxide is added favor the increased formation of undesired polymers 
(green oil). 
The known catalysts for the selective hydrogenation of multiply unsaturated 
compounds are generally prepared by impregnation of an inert support with 
an aqueous solution of a palladium salt, a mixture of a palladium salt 
with a promoter salt or by successive separate impregnation with aqueous 
solutions of the substances active as catalyst and/or promoter, subsequent 
drying and calcination at relatively high temperatures. Most of the 
available catalysts are reduced with hydrogen after installation in the 
reactor. 
DE-A 2 107 568 describes a process for purifying hydrocarbons by selective 
hydrogenation. Multiply unsaturated compounds such as methylacetylene and 
propadiene are hydrogenated in the liquid phase in two reaction zones 
connected in series. In the first reaction zone, part of the liquid 
evaporates. The catalyst used is Pd on Al.sub.2 O.sub.3. 
EP-A-0 653 243 describes supported catalysts which are obtained by 
dissolving palladium nitrate solution, possibly together with silver 
nitrate solution, in a solvent, admixing the solution with a high 
molecular weight sodium polyacrylate and mixing with aluminum oxide as 
support. The composition obtained is shaped, dried and calcined. The 
catalyst is used for the selective hydrogenation of methylacetylene and 
propadiene in a C.sub.3 stream in the liquid phase. 
EP-A-0 532 482 describes a process for the selective hydrogenation of 
butadiene-rich raw C.sub.4 fractions. The selective hydrogenation of 
butadiene to butenes is carried out in the liquid phase over fixed-bed 
supported palladium catalysts. The hydrogenation is carried out in two 
reaction zones connected in series. 
DE-C-31 19 850 describes a process for the selective hydrogenation of 
butadiene in a C.sub.4 fraction. The liquid-phase hydrogenation is carried 
out using 0.3% by weight of palladium on an aluminum oxide support in the 
form of spheres having a diameter of 2 mm. 
EP-B-0 288 362 describes a process for the isomerization of 1-butene to 
2-butenes in a C.sub.4 hydrocarbon fraction comprising butadiene and 
sulfur-containing compounds. The hydrocarbon fraction is passed over a 
first bed of a catalyst comprising palladium and gold and/or platinum. The 
stream is then passed over a second catalyst bed comprising palladium 
deposited on aluminum oxide or on silicon dioxide. 
U.S. Pat. No. 4,260,840 describes a process for purifying a stream 
comprising 1-butene. Here, butadiene is selectively hydrogenated to butene 
in a C.sub.4 stream containing at least 30% by weight of 1-butene. As 
supported catalyst, use was made of Pd/Cr on aluminum oxide in a packed 
catalyst bed. 
U.S. Pat. No. 5,475,173 describes a process for the selective hydrogenation 
of 1,3-butadiene. The catalyst comprises palladium and silver on Al.sub.2 
O.sub.3 and also an alkali metal fluoride. 
EP-B-0 064 301 describes a catalyst for the selective hydrogenation of 
acetylene. The catalyst comprises from 0.01 to 0.025% by weight of 
palladium and from 2 to 6 times this amount of silver on an alpha-Al.sub.2 
O.sub.3 support having a surface area of from 3 to 7 m.sup.2 /g. The 
catalyst has a low CO-sensitivity and a long operating life. 
EP-B-0 089 252 describes a catalyst for the selective hydrogenation of 
acetylenic hydrocarbons. The catalyst comprises from 0.03 to 1% by weight 
of palladium and from 0.003 to 0.3% by weight of gold on an Al.sub.2 
O.sub.3 support. 
U.S. Pat. No. 4,839,329 describes a process for preparing a palladium 
catalyst on a titanium dioxide support. The palladium content is from 0.01 
to 0.2% by weight. The catalyst is suitable for the selective 
hydrogenation of acetylene to ethene. 
DE-C 1 284 403 describes a process for preparing palladium-heavy 
metal-alumina catalysts for the removal of acetylenes and diolefins from 
gas mixtures comprising predominantly monoolefins by selective 
hydrogenation. Pd/Cr on alumina-containing supports is used for removing 
methylacetylene and propadiene. 
DE-C 1 299 629 describes a process for removing acetylenes from gas 
mixtures comprising predominantly olefins by selective hydrogenation. A 
Pd/Cr catalyst on alumina is likewise used for the gas-phase hydrogenation 
of propadiene and methylacetylene. 
The known supported catalysts have the usual disadvantages of 
oxide-supported catalysts. They display abrasion, are sensitive to 
mechanical stress in the case of pressure pulses or the occurrence of a 
pressure drop over the catalyst bed and are unpleasant to handle when 
installing or removing fresh or spent catalyst. 
Catalysis Today, 24 (1995), pages 181-187 describes the use of an 
.alpha.-Al.sub.2 O.sub.3 monolith having a wall thickness of 0.2 mm and a 
cell density of 110 cells/cm.sup.2 for the selective hydrogenation of 
acetylene in the C.sub.2 stream from a steam cracker in the gas and liquid 
phase. 
A disadvantage of the ceramic monoliths is the absence of transverse mixing 
in the individual separated channels and the formation of laminar flows at 
low flow velocities, which leads to poorer selectivities. 
It is an object of the present invention to provide catalysts for the 
selective hydrogenation of multiply unsaturated hydrocarbons in 
hydrocarbon streams. A further object of the present invention is the 
provision of a process for the hydrogenation of multiply unsaturated 
hydrocarbons which avoids the above-described disadvantages of the known 
catalysts. 
We have found that these objects are achieved by a process for the 
hydrogenation of multiply unsaturated C.sub.2 -C.sub.8 hydrocarbons, in 
particular C.sub.2 -C.sub.8 -alkynes and/or C.sub.4 -C.sub.8 -alkynenes 
and/or C.sub.4 -C.sub.8 -alkadienes in fluids comprising these by contact 
with a catalyst packing in the presence of free hydrogen, wherein the 
catalyst packing can be produced by applying at least one substance which 
is active as catalyst and/or promoter to woven meshes knitwear or foils as 
support material. 
The catalysts used according to the present invention have the structure 
described below. 
Support material 
Support materials which can be used for the catalysts employed according to 
the present invention are many foils and woven meshes, as well as knitted 
meshes. According to the present invention it is possible to use woven 
meshes having different types of weave, for example smooth-surface woven 
mesh, twilled mesh, braid-woven mesh, five-shaft satin-woven mesh or other 
special types of weave. Suitable woven wire meshes are, according to one 
embodiment of the invention, meshes made of weavable metal wires such as 
iron, spring steel, brass, phosphor bronze, pure nickel, Monel metal, 
aluminum, silver, nickel silver, nickel, Nichrome, chromium steel, 
stainless, acid-resistant and high-temperature-resistant chromium-nickel 
steels and also titanium. The same applies to knitted meshes. 
It is likewise possible to use woven or knitted meshes of inorganic 
materials, for example of Al.sub.2 O.sub.3 and/or SiO.sub.2. 
Synthetic wires and woven meshes made of plastics can also be used 
according to an embodiment of the invention. Examples are polyamides, 
polyesters, polyvinyls, polyolefins such as polyethylene, polypropylene, 
polytetrafluoroethylene and other plastics which can be processed into 
woven or knitted meshes. 
Preferred support materials are metal foils or woven metal meshes, for 
example stainless steels having the material numbers 1.4767, 1.4401, 
2.4610, 1.4765, 1.4847, 1.4301, etc. The designation of these materials by 
the material numbers mentioned is according to the material numbers in the 
"Stahleisenliste", published by the Verein Deutscher Eisenhuttenleute, 8th 
edition, pages 87, 89 and 106, Verlag Stahleisen mbH, Dusseldorf, 1990. 
The material number 1.4767 is also known under the name Kanthal. 
The metal foils and woven metal meshes are particularly suitable since they 
can be roughened by heating the surface before coating with catalytically 
active compounds or promoters. For this purpose, the metallic supports are 
heated at from 400.degree. to 1100.degree. C., preferably from 800.degree. 
to 1000.degree. C., for from 0.5 to 24 hours, preferably from 1 to 10 
hours, in an oxygen-containing atmosphere such as air. According to an 
embodiment of the invention, this pretreatment can be used to control or 
increase the activity of the catalyst. 
Coating of the catalyst support 
According to the present invention, the catalyst supports used according to 
the present invention can be coated by means of various methods with 
catalytically active compounds and promoters. 
According to one embodiment of the invention, the substances which are 
active as catalyst and/or promoter are applied by impregnation of the 
support in bulk, by electrochemical deposition or deposition in the 
presence of a reducing agent (electroless deposition). 
The catalyst mesh or catalyst foil can then, according to an embodiment of 
the invention, be shaped to form monoliths for installation in the 
reactor. According to a further embodiment of the invention, the shaping 
can also be carried out before application of the active substances or 
promoters. 
According to an embodiment of the invention, the catalyst supports which 
can be used according to the present invention, in particular the woven or 
knitted meshes and foils, can be coated with "thin layers" of 
catalytically active compounds and promoters by means of a vacuum vapor 
deposition technique. For the purposes of the present invention, "thin 
layers" are coatings in the thickness range from a few .ANG. (10.sup.-10 
m) to a maximum of 0.5 .mu.m. As vacuum vapor deposition techniques, 
various processes can be employed according to the present invention. 
Examples are thermal vaporization, flash vaporization, cathode atomization 
(sputtering) and the combination of thermal vaporization and cathode 
atomization. Thermal vaporization can here be carried out by means of 
direct or indirect electric heating. 
Vaporization by means of an electron beam can likewise be used according to 
the present invention. For this purpose, the substance to be vaporized is 
heated on the surface in a water-cooled crucible by means of an electron 
beam so strongly that even high-melting metals and dielectrics are 
vaporized. According to an embodiment of the invention, chemical reactions 
can be effected during buildup of the layers by vapor deposition 
techniques by means of targeted additions of suitable amounts of reactive 
gases to the residual gas. A suitable reaction procedure thus enables 
oxides, nitrides or carbides to be produced on the support. 
Using the process of the present invention, the supports, in particular the 
woven or knitted meshes and foils, can be treated with vapor batchwise or 
continuously in a vacuum vapor deposition unit. For example, the vapor 
treatment is carried out by heating the catalytically active component or 
compound to be applied, for example a noble metal, in a vacuum of from 
10.sup.-2 to 10.sup.-10 torr preferably from 10.sup.-4 to 10.sup.-8 torr, 
by means of an electron beam so strongly that the metal is vaporized out 
of the water-cooled crucible and is deposited on the support. The support 
mesh or knitwear is advantageously arranged such that as great as possible 
a part of the vapor stream condenses on the support. The meshes or 
knitwear can here be coated continuously by means of a winding machine. 
According to the present invention, preference is given to continuous 
sputtering in an air-to-air unit. 
Suitable parameters and conditions for the vacuum vapor deposition 
techniques may be found, for example, in "Handbook of Thin Film 
Technology", Maissel and Glang, McGraw Hill, New York, 1970, "Thin Film 
Processes" by J. L. Vossen and B. Kern, Academic Press, New York, and also 
EP-A 0 198 435. EP-A-0 198 435 discloses the production of a catalyst mesh 
packet by vapor deposition of platinum or platinum and rhodium onto 
stainless steel mesh. 
In the production of the catalyst according to the present invention by 
vacuum vapor deposition techniques, polycrystalline particles which are as 
disordered and disrupted as possible should be produced on the support and 
the predominant proportion of the atoms of the particles should be on the 
surface. The vacuum vapor deposition technique employed here is thus 
different from the known vapor deposition techniques in the optical and 
electrical industries in which a high purity of the support and 
vapor-deposited materials has to be ensured and a predetermined 
condensation temperature on the support as well as a particular vapor 
deposition rate has to be set. 
In the process of the present invention, it is possible for one or more 
catalytically active compounds or promoters to be vapor-deposited. 
According to one embodiment of the invention, the coatings of catalytically 
active substance are preferably in the thickness range from 0.2 nm to 100 
nm, particularly preferably from 0.5 nm to 20 nm, in particular from 3 to 
7 nm. 
According to an embodiment of the invention, catalytically active compounds 
used are the elements of transition group VIII of the Periodic Table of 
the Elements, preferably nickel, palladium and/or platinum, in particular 
palladium. Promoters can be present according to one embodiment of the 
invention and can be selected according to the present invention from, for 
example, the elements of main groups III, IV, V and VI and also transition 
groups I, II, III, VI and VII of the Periodic Table of the Elements. 
The promoter which is used according to one embodiment of the invention is 
preferably selected from the group consisting of copper, silver, gold, 
zinc, chromium, cadmium, lead, bismuth, tin, antimony, indium, gallium, 
germanium, tungsten or mixtures thereof, particularly preferably silver, 
indium and germanium, copper, gold, zinc, chromium, cadmium, lead, 
bismuth, tin, antimony. The layer thickness of the promoter or promoters 
used according to one embodiment of the invention is from 0.1 to 20 nm, 
preferably from 0.1 to 10 nm, in particular from 0.5 to 3 nm. 
Before the application of the catalytically active substance and/or the 
promoter, the support can be modified by vapor deposition of a layer of is 
an oxidizable metal and subsequent oxidation to form an oxide layer. 
According to one embodiment of the invention, the oxidizable metal used is 
magnesium, aluminum, silicon, titanium, zirconium, tin or germanium or a 
mixture thereof. The thickness of such an oxide layer is, according to the 
present invention, in the range from 0.5 to 200 nm, preferably from 0.5 to 
50 nm. 
The coated support material can be heat-treated after coating, for example 
a palladium-coated support material at from 200.degree. to 800.degree. C., 
preferably from 300.degree. to 700.degree. C., for from 0.5 to 2 hours. 
After the catalyst has been produced, it can, if desired or necessary, be 
reduced with hydrogen at from 20.degree. to 250.degree. C., preferably 
from 100.degree. to 200.degree. C. This reduction can also be carried out 
in the reactor itself, which is preferred. 
According to an embodiment of the invention, the catalysts can be built up 
systematically, for example in a vapor deposition unit using a plurality 
of different vaporization sources. Thus, for example, an oxide layer or, 
by reactive vapor deposition, a bonding layer can first be applied to the 
support. Catalytically active components and promoters can be 
vapor-deposited on this base layer in a plurality of alternating layers. 
Introducing a reactive gas into the receptacle during vapor deposition 
enables promoter layers of oxides and other compounds to be produced. 
Heat-treatment steps can also be carried out in between or subsequently. 
The substance or substances active as catalyst and/or promotor can also be 
applied by impregnation. 
The catalysts produced by vapor deposition according to the present 
invention, in particular catalyst meshes, catalyst knitwear and catalyst 
foils, have very good adhesion of the catalytically active compounds or 
promoters. They can therefore be shaped, cut and, for example, processed 
into monolithic catalyst elements without the catalytically active 
compounds or promoters being detached. Catalyst packings of any shape for 
a reactor, eg. flow-through reactor, a reaction column or distillation 
column can be produced from the catalyst meshes, catalyst knitwear and 
catalyst foils of the present invention. It is possible to produce 
catalyst packing elements having different geometries, as are known from 
distillation and extraction technology. Examples of advantageous catalyst 
packing geometries according to the present invention which offer the 
advantage of a low pressure drop in operation are those of the structural 
type Montz A 3 and Sulzer BX, DX and EX. An example of a catalyst geometry 
according to the present invention made of catalyst foils or expanded 
metal catalyst foils are those of the type Montz BSH. 
The amount of catalyst, in particular amount of catalyst mesh, catalyst 
knitwear or amount of catalyst foil, processed per unit volume can be 
controlled within a wide range, whereby a different size of the openings 
or channel widths in the catalyst mesh or catalysr knitwear or in the 
catalyst foil is obtained. Appropriate selection of the amount of catalyst 
mesh, catalyst knitwear or catalyst foil per unit volume enables the 
maximum pressure drop in the reactor, eg. flow-through or distillation 
reactor, to be set and thus enables the catalyst to be matched to 
experimentally determined requirements. 
The catalyst used according to the present invention preferably has a 
monolithic form as is described, for example, in EP-A-0 564 830. Further 
suitable catalysts are described in EP-A-0 218 124 and EP-A-0 412 415. 
A further advantage of the monolithic catalysts used according to the 
present invention is the good fixability in the reactor bed, so that, for 
example, they can be used very well in hydrogenations in the liquid phase 
in the upflow mode at a high cross-sectional loading. In comparison, in 
the case of conventional catalyst supports there is the danger of 
fluidization in the catalyst bed which can lead to possible abrasion or 
disintegration of the shaped bodies. In gas-phase hydrogenation, the 
catalyst packing is capable of withstanding shock or vibrations. No 
abrasion occurs. 
Hydrogenation 
The above-described catalysts are used according to the present invention 
in processes for the hydrogenation, in particular selective hydrogenation, 
of multiply unsaturated C.sub.2 -C.sub.8 -hydrocarbons in fluids 
comprising these. The multiply unsaturated hydrocarbons can be, for 
example, C.sub.2 -C.sub.8 -alkynes, C.sub.4 -C.sub.8 -alkynenes, C.sub.4 
-C.sub.8 -alkadienes or mixtures of these. They are preferably unsaturated 
C.sub.2 -C.sub.6 -hydrocarbons, in particular C.sub.2 -C.sub.4 
-hydrocarbons. 
According to an embodiment of the invention, these multiply unsaturated 
hydrocarbons are present in C.sub.2, C.sub.3, C.sub.4, C.sub.5 or C.sub.6 
streams, preferably in streams from a steam cracker or catalytic cracker. 
These streams generally comprise, as described above, more or less large 
amounts of the corresponding multiply unsaturated C.sub.2 -C.sub.6 
-hydrocarbons. 
By using the catalysts of the present invention, these compounds can be 
converted into the corresponding monounsaturated hydrocarbons with high 
selectivity and in high yield. 
The selective hydrogenations are, according to the present invention, 
carried out either adiabatically or isothermally in the gas or liquid 
phase. The number of reactors depends on the amount of compounds to be 
hydrogenated in the gas stream or liquid stream. For example, an 
adiabatically operated reactor suffices for contents below 1% by weight in 
gas-phase hydrogenations, with the hydrogen/multiply unsaturated 
hydrocarbon ratio being from about 1.8 to 2. If the content of multiply 
unsaturated compounds is higher, the hydrogenation is carried out in two 
or more reactors connected in series. In this case, the hydrogen is fed in 
before each reactor. 
The hydrogenation of a C.sub.3 stream in the gas phase is usually carried 
out in three reactors connected in series, with a conversion of from 60 to 
70% being achieved in the first reactor and a conversion of from 30 to 40% 
being achieved in the second reactor. The remaining conversion is achieved 
in the third reactor, or the third reactor serves as a safety reactor. 
In the case of hydrogenation in the liquid phase, an adiabatically operated 
reactor without recirculation suffices for contents of multiply 
unsaturated hydrocarbons of up to 3.3% by weight. At a hydrogen/multiply 
unsaturated hydrocarbon ratio of from about 1 to 1.5, this gives a 
depletion down to from 500 to 1000 ppm in the output, which corresponds to 
a conversion of from 95 to 99%. If the content of multiply unsaturated 
hydrocarbons is higher, recirculation is generally necessary. If the 
content of multiply unsaturated hydrocarbons in the output is to be 
reduced to less than 10 ppm, the hydrogenation is generally carried out in 
two reactors connected in series, with the hydrogen being fed in before 
each reactor as described above. At a hydrogen/unsaturated hydrocarbon 
ratio of from about 4 to 8, a total conversion of more than 99.9% is 
achieved in the second reactor. 
In the hydrogenation of C.sub.2 streams having acetylene contents of more 
than 2% by weight, the hydrogenation is usually carried out in one 
isothermal reactor and one or two adiabatic reactors connected to the 
isothermal reactors. 
In the liquid-phase hydrogenation of a C.sub.4 stream with high content of 
butadiene, one or two stages are provided depending on the desired 
butadiene depletion. Above a depletion factor of about 200, a two-stage 
process is generally preferred. Thus, for example, the selective 
hydrogenation of a raw C.sub.4 stream from a steam cracker containing 
about 45% by weight of butadiene is carried out in two stages to a 
residual butadiene content of less than 10 ppm. 
It is as well possible to remove low contents of butadiene selectively in 
as so called remainder hydrogenation. In this case a one step process with 
depletion factors of more than 1000 is accessible. For example the 
hydrogenation of 0,5% by weight of butadiene to values below 10 ppm is 
performed in a one step process, wherein at the same time a maximum of 
butene-1 present can be retained. 
According to one embodiment of the invention, the hydrogenation is carried 
out in the gas phase. In particular, the hydrogenation of C.sub.2 and/or 
C.sub.3 streams is carried out in the gas phase. Examples of reactors 
which can be used are tube reactors and shaft reactors as well as 
tube-bundle reactors. 
According to one embodiment of the invention, a plurality of tube reactors 
can be connected in series. Here, according to one embodiment of the 
invention, the hydrogen is fed in before each reactor. For a further 
description of reactors which are suitable according to the present 
invention, reference is made to the introduction. 
The selective hydrogenation in the gas phase is, according to one 
embodiment of the invention, carried out at pressures of from 5 to 50 bar, 
preferably from 10 to 30 bar, in particular from 15 to 25 bar. According 
to one embodiment of the invention, the space velocities are from 500 to 
8000 m.sup.3 /m.sup.3 h, preferably from 1000 to 5000 m.sup.3 /m.sup.3 h, 
in particular from 2000 to 4000 m.sup.3 /m.sup.3 h. The inlet temperature 
for the hydrogenation is, according to one embodiment of the invention, 
from -20.degree. to 150.degree. C., preferably from 20.degree. to 
120.degree. C., in particular from 20.degree. to 80.degree. C. It is 
possible to use an adiabatically operated or an isothermally operated 
reactor. The hydrogenation can likewise be carried out in a plurality of 
reactors connected in series, these being operated isothermally or 
adiabatically. For example, two adiabatic reactors can follow one 
isothermal reactor, particularly in the hydrogenation of a C.sub.2 stream. 
According to an embodiment of the invention, the hydrogenation is carried 
out in the liquid phase or in a mixed liquid/gas phase with at least 50% 
by weight of the hydrocarbon stream in the liquid phase. Here, according 
to an embodiment of the invention, the hydrogenation can be carried out in 
the downflow mode or in the upflow mode. In the upflow made the hydrogen 
added can be present as a solution in the liquid phase. Reactors which can 
be used here are, for example, tube reactors or tube-bundle reactors. 
According to one embodiment, the hydrogenation is carried out at a pressure 
of from 5 to 70 bar, preferably from 5 to 40 bar, in particular from 10 to 
30 bar. According to one embodiment of the invention, the space velocity 
is from 1 to 100 m.sup.3 /m.sup.3 h, preferably from 2 to 40 m.sup.3 
/m.sup.3 h, in particular from 2 to 20 m.sup.3 /m.sup.3 h. The inlet 
temperature for the hydrogenation is, according to one embodiment of the 
invention, from -10.degree. to 150.degree. C., preferably from 0.degree. 
to 120.degree. C., in particular from 0.degree. to 90.degree. C. In order 
to ensure the formation of a liquid phase, it is necessary to select 
suitable temperature and pressure parameters which are dependent on the 
mixture of substances used in each case. 
According to an embodiment of the invention, the hydrogenation is carried 
out in a catalytic distillation process. In this process, the 
hydrogenation as described above is combined with simultaneous 
distillation or rectification over the catalyst packing. 
In such a process, the hydrogenation and a distillation take place 
simultaneously or immediately after one another. At least one component of 
the reaction mixture is distilled from the hydrogenation mixture after the 
hydrogenation. The term "catalytic distillation" refers to a chemical 
reaction, here a hydrogenation, which is combined with a distillation or 
rectification in a suitable apparatus. As reactor for the catalytic 
distillation, it is possible to use any suitable distillation apparatus in 
which the catalyst packing can be installed in the distillation part. This 
is possible, for example, by installation of the catalyst packing in a 
distillation column in the distillation apparatus. 
The reaction mixture, ie. the hydrocarbon stream, is introduced into the 
distillation apparatus at a suitable point, according to one embodiment 
into the bottom of the distillation apparatus. This is particularly 
advantageous in the hydrogenation of a C.sub.3, C.sub.4, C.sub.5 or 
C.sub.6 stream. The hydrogenated components and the alkenes are here taken 
off at the top of the distillation apparatus. 
Preferably the hydrogenation is proceeds selectivetly and essentially no 
hydrogenation of alkenes to alkanes occnos. 
The invention is illustrated by the following examples. 
In the performance tests for hydrogenation in the gas phase, the monolithic 
catalysts were used in an unpressurized laboratory apparatus or in a pilot 
part apparatus under increased pressures. The temperatures of the gas 
mixture entering the hydrogenation zone are generally from 15.degree. to 
about 120.degree. C., preferably from 25.degree. to 90.degree. C. The 
volume ratio of hydrogen to the multiply unsaturated hydrocarbons is 
generally from 0.5:1 to 2.5:1, in the C.sub.2 hydrogenation preferably 
from 1.1:1 to 2:1, in particular from 1.2:1 to 1.8:1, and in the first 
stage of the C.sub.3 hydrogenation from 0.5:1 to 0.8:1. 
In the following, proportions by volume of gas are proportions by volume at 
STP.

EXAMPLE 1 
Plain-woven wire mesh made of material No. 1.4301 and having a mesh opening 
of 0.125 mm and a wire diameter of 0.1 mm was heated in air at 800.degree. 
C. for 3 hours. After cooling, the support mesh which had been pretreated 
in this way had first 6 nm of Pd and then 1 mm of Ag vapor-deposited on 
both sides in an electron beam vapor deposition unit at a pressure of from 
1 to 3.times.10.sup.-6 torr. The thickness of the layers was measured by 
means of a crystal oscillator and the vapor deposition rate was controlled 
using the crystal oscillator. The amount of palladium deposited was 138 
mg/m.sup.2 and the amount of silver was 19.5 mg/m.sup.2. The catalyst mesh 
thus produced was fabricated into 3 monoliths having a height of 90 mm and 
a diameter of 18.6 mm. In the middle of the monoliths there was a 
thermocouple hole having a diameter of 4 mm. To produce the monoliths, 
mesh strips having a width of 92 mm and a length of 37.5 cm were cut and 
one of these was corrugated by means of a toothed roller (modulus 0.5 mm). 
This corrugated mesh was laid together with the smooth mesh and wound 
around a 4 mm thick metal rod. This gave a monolithic catalyst which was 
strengthened by point welding at the outer edge. 
EXAMPLE 2 
Gas-phase hydrogenation of a C.sub.3 stream under pressure 
Three monoliths produced as described in Example 1 and having a total 
surface area of 4219 cm.sup.2 were installed in a reactor for the test on 
the gas-phase hydrogenation of methylacetylene and propadiene in a C.sub.3 
stream from a steam cracker. A multiple thermocouple was introduced axially 
into the 4 mm wide thermocouple hole. 
The process conditions were set according to the conditions in a first 
stage of the usually 3-stage selective hydrogenation of methylacetylene 
and propadiene in the C.sub.3 stream. 
The reactor had a diameter of 18.6 mm and a length of 2 m. The catalyst 
monolith had a height of 27 cm and a volume of 70 ml. 
After flushing with nitrogen and hydrogen at 120.degree. C., 660 g/h of a 
gas mixture composed of 6.8% of propane, 1.7% of propadiene and 2.2% of 
methylacetylene in propylene were mixed with differing amounts of hydrogen 
and passed over the catalyst at an inlet temperature of 50.degree. C. and a 
pressure of 10 bar. The compositions of the reaction product are summarized 
in the table below. 
TABLE 1 
__________________________________________________________________________ 
Propad- Convers- 
H.sub.2 / 
Propane 
Propene 
iene Propine 
C6+ ion S(prope- 
H.sub.2 
MA % by 
% by 
% by % by 
% by 
(MAPD) 
ne) 
I/h! 
PD volume 
volume 
volume 
volume 
volume 
%! %! 
__________________________________________________________________________ 
8.1 0.5 
6.848 
91.52 
0.737 
0.67 
0.22 
64 90 
8.9 0.55 
6.895 
91.61 
0.653 
0.594 
0.244 
68 88 
9.7 0.6 
6.962 
91.67 
0.567 
0.53 
0.262 
72 86 
11.3 
0.7 
7.161 
91.72 
0.439 
0.439 
0.268 
78 81 
__________________________________________________________________________ 
MAPD is the mixture of multiply unsaturated hydrocarbons, namely 
methylacetylene and propadiene. The ratio of hydrogen to MAPD is the 
volume ratio. S is the selectivity based on propene. 
Under the conditions of the first hydrogenation stage, the catalyst has 
very high selectivities. 
EXAMPLE 3 
Plain-woven wire mesh made of material No. 1.4767 and having a mesh opening 
of 0.18 mm and a wire diameter of 0.112 mm was heated in air at 900.degree. 
C. for 5 hours. After cooling, the support mesh which had been pretreated 
in this way had first 92 mg of Pd/m.sup.2 and then 26.4 mg of Zn/m.sup.2 
vapor deposited under the same conditions on both sides in an electron 
beam vapor deposition unit at a pressure of 1.times.10.sup.-6 torr. As 
described in Example 1, a wound 126 cm.sup.3 monolith was produced from 
the catalyst mesh thus obtained. 
EXAMPLE 4 
Pressureless C.sub.2 gas-phase hydrogenation 
The catalyst monolith obtained as described in Example 3 was installed in a 
tube reactor as described in Example 2. The test of the catalyst was 
carried out under atmospheric pressure using a gas mixture of 1% by volume 
of acetylene, 2% by volume of hydrogen and 97% by volume of ethylene at a 
space velocity over the catalyst of 3000 m.sup.3 /m.sup.3 (cat) h. At 
82.degree. C., an acetylene conversion of 70% was achieved at a 
selectivity to ethylene of 97%. Under otherwise identical reaction 
conditions, a commercial supported catalyst containing 0.02% by weight of 
Pd and 0.01 % by weight of Zn gave a selectivity to ethylene of only 62% 
at a conversion of 70%. 
EXAMPLE 5 
The support material used was the material described in Example 3 which was 
pretreated by heating in air at 900.degree. C. and subsequently had 138 
mg/m.sup.2 of palladium vapor-deposited onto it using a method similar to 
Example 3. Rolling together one corrugated and one smooth strip of 
catalyst mesh having a width of 10 cm produced a monolith having a 5 mm 
thermocouple hole. The resulting monolith had a volume of 71.6 cm.sup.3 
and comprised 15.25 dm.sup.2 of catalyst mesh. 
EXAMPLE 6 
Gas-phase hydrogenation of a C.sub.2 stream under a pressure of 20 bar 
The catalyst monolith produced as described in Example 5 was installed in a 
tube reactor as described in Example 2. After flushing with nitrogen, the 
catalyst was reduced with 10 l/h of hydrogen for 3 hours at 150.degree. C. 
At an inlet temperature of 82.degree. C., 160 l/h of a gas mixture 
comprising 98.824% by volume of ethylene and 1.145% by volume of 
acetylene, which had been mixed with 1.46% by volume of hydrogen, were 
then passed over the catalyst. The reaction product comprised 99.394% by 
volume of ethylene, 0.486% by volume of ethane and 0.01% by volume of 
acetylene (conversion 99.1%, selectivity 58%). On increasing the hydrogen 
content to 1.67% by volume, the ethane content in the reaction product 
rose to 0.678% by volume. Acetylene could then no longer be detected 
(conversion 100%, selectivity 43%). 
Addition of 1.5 ppm of carbon monoxide enabled the selectivity to be 
increased further. At an inlet temperature of 84.degree. C., an 
acetylene-free reaction product was obtained. The ethylene content was 
99.419% by volume, the ethane content was 0.442% by volume. 
EXAMPLE 7 
Using the method described in Example 1, 138 mg of Pd/m.sup.2 and then 19.5 
mg of Ag/m.sup.2 were vapor-deposited on a supported mesh made of material 
No. 1.4767 which had been pretreated as described in Example 1 by heating 
in air at 900.degree. C. The catalyst mesh was subsequently fabricated 
into a monolith having a volume of 126 cm.sup.3. 
EXAMPLE 8 
Pressureless gas-phase hydrogenation of a C.sub.2 stream 
The catalyst produced as described in Example 1 was used as described in 
Example 4 for the selective hydrogenation of acetylene. At a conversion of 
70%, a selectivity to ethylene or 91% was achieved. 
EXAMPLE 9 
Liquid-phase hydrogenation of C.sub.3 streams 
The catalyst produced as described in Example 1 was used for the 
liquid-phase hydrogenation of methylacetylene and propadiene in a C.sub.3 
stream from a steam cracker. The process conditions were selected 
corresponding to the conditions of a first stage of the usually 2-stage 
hydrogenation of the C.sub.3 stream. 
3 of the monoliths produced as described in Example 1 and having a total 
surface area of 4219 cm.sup.2 were installed in an adiabatically operated 
tube reactor having a diameter of 20 mm. A multiple thermocouple was 
introduced axially into the 4 mm wide thermocouple hole. To ensure good 
wetting of the catalyst, as is ensured in industrial reactors by the high 
cross-sectional loading, the upflow mode was employed. After flushing with 
nitrogen and hydrogen at 120.degree. C., 520 g/h of a C.sub.3 stream from a 
steam cracker composed of 6.8% by volume of propane, 1.7% by volume of 
propadiene and 2.2% by volume of methylacetylene in propylene and 13 
standard l/h of hydrogen were passed over is the catalyst at an inlet 
temperature of 10.degree. C. and a pressure of 23 bar. The reaction 
product comprised 7.3% by volume of propane and 0.3% by volume of unknowns 
(oligomers) in propylene. This corresponds to a selectivity to propylene of 
80% at a conversion of more than 99.9%. The results are summarized in Table 
COMATIVE EXAMPLE 1 
The catalyst produced as described in EP-A-0 653 243 was used as a 
comparative catalyst in the liquid-phase hydrogenation of methylacetylene 
and propadiene in the C.sub.3 stream from a steam cracker. The process 
conditions were selected as in Example 9. 70 ml of the catalyst were 
installed in the adiabatically operated tube reactor. After flushing with 
nitrogen and hydrogen at 120.degree. C., 520 g/h of a C.sub.3 stream from 
a steam cracker composed of 5.1% by volume of propane, 1.8% by volume of 
propadiene, 2.3% by volume of methylacetylene in propylene and 13 standard 
l/h of hydrogen were passed over the catalyst at an inlet temperature of 
10.degree. C. and a pressure of 22 bar. The reaction product comprised 
5.5% by volume of propane and 0.5% by volume of unknowns in propylene. The 
unknowns are oligomers formed. This corresponds to a selectivity to 
propylene of 78% at a conversion of greater than 99.9%. The results are 
summarized in Table 2. 
COMATIVE EXAMPLE 2 
The catalyst LD 265 described in Chem. Eng. Prog., 70 (1974), 74-80 was 
employed as a comparative catalyst for the liquid-phase hydrogenation of 
methylacetylene and propadiene in a C.sub.3 stream from a steam cracker. 
The process conditions were selected as in Comparative Example 1, but the 
stream contained 8% of propane, 1.7% by volume of propadiene and 2.1% by 
volume of methylacetylene. The reaction product comprised 8.5% by volume 
of propane and 0.7% by volume of unknowns in propylene. This corresponds 
to a selectivity to propylene of 69% at a conversion of more than 99.9%. 
The results are summarized in Table 2. 
TABLE 2 
______________________________________ 
Comparative Comparative 
Ex. 1 Ex. 2 
(0.3% Pd, 0.4 
(0.3% Pd, 0.7 
Example 9 
Catalyst kg/l) kg/l) Pd/Ag-Catalyst 
______________________________________ 
whsv kg/l! 
ca. 6.5 ca. 6.5 ca. 6.5 
Pressure bar! 
22 22 22 
H.sub.2 /MAPD calc. 
1.08 1.1 1.1 
mol/mol! 
T.sub.in .degree.C.! 
10 10 10 
MAPD ppm! 
&lt;10 &lt;10 &lt;10 
S(propene) %! 
78 69 80 
Conversion %! 
&gt;99.9 &gt;99.9 &gt;99.9 
.DELTA. Propane %! 
0.5 0.5 0.5 
.DELTA. Unknowns %! 
0.5 0.7 0.3 
______________________________________ 
In the table, whsv is the weight hourly space velocity in kg/l. MAPD is the 
amount of multiply unsaturated hydrocarbons, namely methylacetylene and 
propadiene. The indicated ratios of H.sub.2 /MAPD were calculated from the 
amounts of H.sub.2 consumed in the reaction. 
The table shows an increase in the selectivity from Comparative Example 2 
through Comparative Example 1 to Example 9. Although the thin-layer 
catalyst employed in Example 9 contains only 28 mg of Pd and 4 mg of Ag in 
the amount of catalyst used and, for example, the catalyst in Comparative 
Example 2 contains 240 mg of Pd, it has a comparable activity and a higher 
selectivity. The formation of oligomers summarized as unknowns is lowest 
for the catalyst of the invention used in Example 9. 
EXAMPLE 10 
To produce the catalyst, a plain-woven wire mesh made of material No. 
1.4767 and having a mesh opening of 0.18 mm and a wire diameter of 0.112 
mm was heated in air at 900.degree. C. for 5 hours. After cooling, 138 mg 
of Pd/m.sup.2 of mesh were deposited on both sides of the support material 
at a pressure of 1.times.10.sup.-6 torr. Monoliths were subsequently 
produced from this catalyst mesh. For this purpose, a 20 cm wide mesh 
strip was corrugated by means of a toothed roller (modulus 0.5 mm) and 
together with a smooth mesh was rolled up around a metal rod having a 
diameter of 4.5 mm to give a roll. The roll was strengthened by point 
welding at the outer edge and the metal rod was removed to leave the 
thermocouple hole. The monolithic catalyst thus obtained had a diameter of 
16 mm and a height of 20 cm. The amount of catalyst mesh in a monolith was 
940 cm.sup.2 and 5 monoliths were installed in the hydrogenation reactor. 
EXAMPLE 11 
Liquid-phase hydrogenation of raw C.sub.4 fraction from a steam cracker. 
The selective hydrogenation of a raw C.sub.4 fraction was carried out over 
the catalyst from Example 10 in a fixed-bed reactor of a pilot plant unit 
which was fitted with a separator and a liquid circuit. The fixed-bed 
reactor was able to be heated by means of electric heating and had a 
diameter of 16 mm and a length of 2 m. The starting material was metered 
into the circulating stream by means of a pump and mixed with the 
necessary hydrogen at a mixing point. The selective hydrogenation was 
carried out in a fixed bed comprising the monolithic catalyst described in 
Example 10. The reaction mixture subsequently went to a separator in which 
the gas and liquid phases were separated. The major part of the liquid 
phase was circulated. A smaller part corresponding to the amount of 
starting material was continuously taken from the system and analyzed by 
gas chromatography. 
Before commencement of the experiment, the installed monolithic catalyst 
was reduced with hydrogen at 120.degree. C. and 5 bar pressure for 12 
hours. The unit was subsequently run up using hydrogenated C.sub.4 
fraction and hydrogen. The results of the experiment on the selected 
hydrogenation are summarized in Table 3 below. 
TABLE 3 
______________________________________ 
Starting 
Pd catalyst from 
ma- Example 10 
terial Hydrogenation product 
______________________________________ 
Space velocity m.sup.3 /m.sup.3 h! 
9.0 9.0 
Recycle/Feed 8.2 8.2 
T.sub.in .degree.C.! 60 60 
p bar! 17.7 18.3 
Ratio of H.sub.2 /(butadiene + buten- 
0.98 1.02 
yne + butyne) 
Butadiene + butenyne + Butyne % 
34.9 1.8 0.5 
by weight! 
1-butene % by weight! 
14.2 40.3 39.5 
2-trans-butene % by weight! 
4.5 17.6 18.6 
2-cis-butene % by weight! 
3.3 5.7 6.2 
i-butene % by weight! 
23.6 23.6 23.6 
i-butane % by weight! 
3.0 3.0 3.0 
n-butane % by weight! 
7.2 7.7 8.3 
C.sub.5 hydrocarbons % by weight! 
0.3 0.3 0.3 
Conversions %! 95.9 98.9 
Total butene selectivity %! 
98.8 97.5 
______________________________________ 
The catalyst displayed a very high activity. High conversions could be 
achieved even at high space velocities. Even in a hydrogenation to a 
residual butadiene content of 1.8% by weight, the hydrogenation to 
n-butane was only 0.5% by weight. No hydrogenation of the i-butene took 
place. 
EXAMPLE 12 
To produce the catalyst, plain-woven wire mesh made of material No. 1.4767 
and having a mesh opening of 0.18 mm and a wire diameter of 0.112 mm was 
heated in air at 1000.degree. C. for 5 hours. After cooling, 92 mg of 
Pd/m.sup.2 were vapor-deposited on both sides of the support material at a 
pressure of 1.times.10.sup.-6 torr. To increase the selectivity, the Pd 
catalyst mesh was subsequently doped with 0.5 nm of germanium by vapor 
deposition. The thickness of the germanium doping layer was measured 
during the vapor deposition procedure using a crystal oscillator. 5 
monoliths were fabricated as described in Example 10 from the catalyst 
mesh thus obtained and these were installed in the hydrogenation reactor. 
EXAMPLE 13 
Liquid-phase hydrogenation of raw C.sub.4 fraction from a steam cracker 
The catalyst described in Example 12 was likewise used in the unit 
described in Example 11. Before commencement of the experiment, the 
catalyst was reduced with hydrogen at 120.degree. C. and 5 bar pressure 
for 12 hours in a similar way to Example 11. The unit was subsequently run 
up using hydrogenated C.sub.4 fraction and hydrogen. The results of the 
experiment on the selective hydrogenation are summarized in Table 4 below. 
TABLE 4 
______________________________________ 
Starting ma- 
Pd/Ge catalyst 
terial from Example 12 
______________________________________ 
Space velocity m.sup.3 /m.sup.3 h! 
9.0 
Recycle/Feed 8.2 
T.sub.in .degree.C.! 60 
p bar! 17.2 
Ratio of H.sub.2 /(butadiene + buten- 
0.97 
yne + butyne) 
Butadiene + butenyne + Butyne % 
46.4 2.4 
by weight! 
1-butene % by weight! 
15.2 42.5 
2-trans-butene % by weight! 
5.1 18.9 
2-cis-butene % by weight! 
3.8 6.3 
i-butene % by weight! 
23.9 23.9 
i-butane % by weight! 
1.0 1.0 
n-butane % by weight! 
4.4 4.8 
C.sub.5 hydrocarbons % by weight! 
0.2 0.2 
Conversion %! 94.8 
Total butene selectivity %! 
99.1 
______________________________________ 
The catalyst had a very high activity. In use, it enabled high space 
velocities to be employed while at the same time achieving a high 
conversion. Compared with the catalyst from Example 11, the total butene 
selectivity is somewhat improved and is above 99%. No hydrogenation of the 
i-butene took place. 
EXAMPLE 14 
Using a method similar to Example 12, metal mesh made of material No. 
1.4767 was heated in air at 1000.degree. C. for 5 hours. After cooling, 
the support mesh was coated with 50 nm of Mg in the vacuum coating unit 
described. The thickness of the layer was measured during the vapor 
deposition procedure using a crystal oscillator. The mesh was subsequently 
heated to 300.degree. C. over a period of 60 minutes and left at this 
temperature in air for 30 minutes. After again being installed in the 
coating unit, it was coated with 6 nm of Pd at 1.times.10.sup.-6 torr. 5 
monoliths were fabricated from the catalyst mesh thus obtained using a 
method similar to Example 10 and these were installed in the hydrogenation 
reactor. 
EXAMPLE 15 
Liquid-phase hydrogenation of raw C.sub.4 fraction from a steam cracker 
The catalyst produced as described in Example 14 was likewise tested in the 
unit described in Example 11. Before commencement of the experiment, the 
catalyst was reduced with hydrogen at 100.degree. C. and 5 bar pressure 
for 12 hours in a manner similar to Example 11. The unit was subsequently 
run up using hydrogenated C.sub.4 fraction and hydrogen. The results of 
the experiment on the selective hydrogenation are summarized in Table 5 
below. 
TABLE 5 
______________________________________ 
Starting ma- 
Pd/MgO catalyst 
terial from Example 14 
______________________________________ 
Space velocity m.sup.3 /m.sup.3 h! 
9.0 
Recycle/Feed 8.2 
T.sub.in .degree.C.! 60 
p bar! 16.3 
Ratio of H.sub.2 /(butadiene + buten- 
0.97 
yne + butyne) 
Butadiene + butenyne + Butyne % 
44.1 2.9 
by weight! 
1-butene % by weight! 
14.2 39.7 
2-trans-butene % by weight! 
4.6 17.4 
2-cis-butene % by weight! 
3.3 5.8 
i-butene % by weight! 
23.6 23.9 
i-butane % by weight! 
2.9 2.9 
n-butane % by weight! 
7.1 7.5 
C.sub.5 hydrocarbons % by weight! 
0.2 0.2 
Conversion %! 93.4 
Total butene selectivity %! 
99.0 
______________________________________ 
The catalyst likewise displayed a high activity and enabled a high space 
velocity to be employed while at the same time achieving a high 
conversion. The performance data are similar to those of the catalyst from 
Example 13. No hydrogenation of the i-butene took place. 
As shown by the examples, the catalysts of the present invention are very 
suitable for the selective hydrogenation of multiply unsaturated 
hydrocarbons. 
Liquid-phase hydrogenation of butadiene-containing raffinate 1 from a steam 
cracker 
COMATIVE EXAMPLE 3 
A Pd,Ag/Al.sub.2 O.sub.3 catalyst produced as described in DE-A-31 19 850, 
Example 3, was used as a comparative catalyst in the liquid-phase 
hydrogenation of butadiene-containing raffinate 1 from a steam cracker. 
The selective hydrogenation of the butadiene was carried out in the pilot 
plant unit described in Example 11. 
Before commencement of the experiment, the Pd,Ag comparative catalyst was 
reduced with hydrogen at 120.degree. C. and 5 bar pressure for 12 hours. 
The pilot plant was subsequently run up using butadiene-containing 
raffinate 1 and hydrogen. The results of this experiment are summarized in 
Table 6. 
EXAMPLE 16 
To produce the catalyst according to the present invention, plain-woven 
wire mesh made of material No. 1.4301 and having a mesh opening of 0.180 
mm and a wire diameter of 0.105 mm was heated in air at 800.degree. C. for 
3 hours. After cooling, the support mesh which had been pretreated in this 
way was coated with 5 nm of Pd and 1 nm of Ag by sputtering in a 
rollcoater. Monoliths were subsequently produced from the catalyst mesh. 
For this purpose, a 20 cm-wide mesh strip was corrugated by means of a 
toothed roller (modulus 0.5 mm) and, using a method similar to Example 10, 
five monoliths having a diameter of 16 mm, a height of 20 cm and an 
internal thermocouple hole having a diameter of 4.5 mm were produced. The 
amount of catalyst mesh for one monolith was 1180 cm.sup.2. The five 
monoliths were finally installed in the hydrogenation reactor which is 
described in Example 11. 
Before commencement of the experiment, the Pd,Ag catalyst according to the 
present invention was reduced with hydrogen at 120.degree. C. and 5 bar 
pressure for 12 hours. The pilot plant was subsequently run up using 
butadiene-containing raffinate 1 and hydrogen. The results of this 
experiment are summarized in Table 6. 
EXAMPLE 17 
To produce the catalyst according to the present invention, plain-woven 
wire mesh made of the material No. 1.4767 and having a mesh opening of is 
0.18 mm and a wire diameter of 0.112 mm was heated in air at 900.degree. 
C. for 5 hours. After cooling, the support mesh which had been pretreated 
in this way had first 4 nm of Pd and then 2 nm of Ag vapor-deposited on 
both sides at a reduced pressure of 1.times.10.sup.-6 torr. The thickness 
of the layers was measured by means of a crystal oscillator and the vapor 
deposition rate was controlled using the crystal oscillator. Monoliths 
were subsequently produced from this catalyst mesh. For this purpose, a 20 
cm-wide mesh strip was corrugated by means of a toothed roller (modulus 0.5 
mm) and, using a method similar to Example 10, five monoliths having a 
diameter of 16 mm, a height of 20 cm and an internal thermocouple hole 
having a diameter of 4.5 mm were produced. The amount of catalyst mesh for 
one monolith was 940 cm.sup.2. The five monoliths were finally installed in 
the hydrogenation reactor which is described in Example 11. 
Before commencement of the experiment, the Pd,Ag catalyst according to the 
present invention was reduced with hydrogen at 120.degree. C. and 5 bar 
pressure for 12 hours. The pilot plant was subsequently run up using 
butadiene-containing raffinate 1 and hydrogen. The results of this 
experiment are summarized in Table 6. 
Table 6 shows a performance comparison for the conventional catalyst from 
Comparative Example 3 and the two catalysts according to the present 
invention from Examples 16 and 17. As can be seen, the catalyst according 
to the present invention from Example 16 gives a 1-butene yield which is 
about 3% higher than that obtained using the comparative catalyst 
described at the same final butadiene content in the hydrogenated product 
of 20 ppm. The advantages of the monolithic catalyst according to the 
present invention from Example 17 are considerably more pronounced, with a 
residual butadiene content in the hydrogenated product of 10 ppm being 
achieved. The 1-butene yield obtained here was over 97%. 
The performance data reveal four significant advantages of the catalyst 
according to the present invention over the comparative catalyst 
described: 
(i) smaller H.sub.2 /butadiene ratio (1.6 instead of 1.9 for the 
comparative catalyst) 
(ii) less overhydrogenation to give n-butane (n-butane formation of 0.4% by 
weight instead of 0.8% by weight for the comparative catalyst) 
(iii) significantly higher 1-butene yield (97.4% instead of 89.2% for the 
comparative catalyst) 
(iv) significantly lower active component content as high activity (12.3 mg 
of active component in the amount of catalyst used instead of 480 mg for 
the comparative catalyst). 
In all examples, no hydrogenation of the i-butene was found. 
EXAMPLE 18 
The conventional comparative catalyst described in Comparative Example 3 
and the catalyst according to the present invention described in Example 
17 were likewise tested under more severe hydrogenation conditions in the 
pilot plant unit described in Example 11. Under these conditions, a 
residual butadiene content in the hydrogenated product of &lt;10 ppm was able 
to be achieved. The results obtained are summarized in Table 7. 
TABLE 7 
______________________________________ 
Conventional 
Pd, Ag/Al.sub.2 O.sub.3 
catalyst from 
Comparative Ex- 
Pd, Ag catalyst 
ample 3 from Example 17 
Hydro- Hydro- 
genation genation 
Feed product Feed product 
______________________________________ 
Space velocity m.sup.3 /m.sup.3 h! 
15 15 
Recycle/feed 1 1 
T.sub.in .degree.C.! 60 60 
p bar! 11.9 11.3 
H.sub.2 /butadiene ratio 
2.7 2.1 
Butadiene % by weight! 
0.43 &lt;0.001 0.54 &lt;0.001 
1-Butene % by weight! 
25.1 20.8 27.2 25.9 
trans-2-Butene % by weight! 
7.9 10.2 8.4 9.0 
cis-2-Butene % by weight! 
5.4 7.1 5.7 6.3 
i-Butene (% by weight! 
42.2 42.2 43.9 43.9 
i-Butane % by weight! 
4.7 4.7 3.0 3.0 
n-Butane % by weight! 
14.0 14.8 11.0 11.7 
C.sub.5 hydrocarbons % by weight! 
0.27 0.2 0.26 0.2 
Conversion %! &gt;99.8 &gt;99.8 
n-Butane formation 0.8 0.7 
% by weight! 
1-Butene yield %! 82.9 95.2 
______________________________________ 
In a hydrogenation to butadiene values of &lt;10 ppm, the catalyst according 
to the present invention likewise shows the abovementioned advantages of a 
small H.sub.2 /butadiene ratio, less overhydrogenation to give n-butane and 
a significantly higher 1-butene yield. As in the previous examples, no 
hydrogenation of i-butene was found in this case. 
TABLE 6 
__________________________________________________________________________ 
Conventional 
Pd, Ag/Al.sub.2 O.sub.3 catalyst 
from Comparative 
Pd, Ag catalyst from 
Pd, Ag catalyst from 
Example 3 Example 16 
Example 17 
Hydrogenation 
Hydrogenation 
Hydrogenation 
Feed 
product 
Feed 
product 
Feed 
product 
__________________________________________________________________________ 
Space velocity m.sup.3 /m.sup.3 h! 
15 15 15 
Recycle/feed 1 1 1 
T.sub.in .degree.C.! 
60 60 60 
p bar! 11.8 11.5 11.3 
H.sub.2 /butadiene ratio 
1.9 1.9 1.6 
Butadiene % by weight! 
0.46 
0.002 0.50 
0.002 0.54 
0.001 
1-Butene % by weight! 
25.0 
22.3 27.7 
25.6 27.2 
26.5 
trans-2-Butene % by weight! 
7.8 
9.1 8.4 
9.4 8.4 
8.9 
cis-2-Butene % by weight! 
5.4 
6.4 5.7 
6.5 5.7 
6.1 
i-Butene % by weight! 
42.9 
42.9 43.6 
43.5 43.9 
43.9 
i-Butane % by weight! 
4.6 
4.6 3.2 
3.2 3.0 
3.0 
n-Butane % by weight! 
13.6 
14.4 10.6 
11.5 11.0 
11.4 
C.sub.5 hydrocarbons % by weight! 
0.24 
0.3 0.3 
0.3 0.26 
0.2 
Conversion %! 99.6 99.6 99.8 
n-Butane formation % by weight! 
0.8 0.9 0.4 
1-Butene yield %! 
89.2 92.4 97.4 
__________________________________________________________________________