GB1596959A

GB1596959A - Selective hydrogenation of c2-minus fractions

Info

Publication number: GB1596959A
Application number: GB23871/78A
Authority: GB
Original assignee: Leuna Werke GmbH
Current assignee: Leuna Werke GmbH
Priority date: 1977-06-01
Filing date: 1978-05-30
Publication date: 1981-09-03
Also published as: DE2819791A1; HU186724B; AT358546B; RO81336A; BG43073A1; DD131644B1; CS206644B1; DD131644A1; DE2819791C2; ATA332378A

Description

(54) SELECTIVE HYDROGENATION OF C2 MINUS FRACTIONS (71) We, VEB LEUNA-WERKE "WALTER ULBRICHT", of 422 Leuna 3, German Democratic Republic, a Corporation organised and existing under the laws of the German Democratic Republic, do hereby declare the invention, for which we pray that a Patent may be granted to us, and the method by which it is to be performed, to be particularly described in and by the following statement: The process for selective hydrogenation of C2-minus fraction relates to a heterogeneously catalysed chemical process, in which, within the scope of the technical production of ethylene from cracked gases of petroleum pyrolysis, the acetylene present in the raw fraction and which causes problems in processing the olefin, is very largely removed, without incurring a substantial loss of ethylene.

The production of ethylene as an intermediate product in the production of plastics has gained enormous technical importance. Preferably, it is obtained by separating and refining the cracked gases from petroleum pyrolysis. Selective hydrogenation of the acetylenic bond has proved to be the almost universal refining method. Selective hydrogenation processes of this kind are required to reduce the acetylene content from 0.3 to 2.0%, normally occurring in the raw fraction, to less than 3 ppm, without substantial quantities of the ethylene, which is present in a large excess, being hydrogenated.

For energy reasons and in order to keep the volume of polymer formation on the catalyst as small as possible, it is generally desired to have as low a working temperature as possible, since, it is known that polymer formation, independently of the particular catalyst system used, is increasingly promoted with rising temperature and inactivation of the catalyst is accelerated. Additionally, the ethylene losses rise extremely rapidly with increasing temperature in the presence of excess hydrogen, even if selectivity-increasing moderators, such as carbon monoxide, are present. Nevertheless, selective hydrogenation of the acetylene in the C2-minus fractions on palladium catalysts, i.e. in a feedstock, which, apart from acetylene, ethylene and ethane, also contains carbon monoxide, methane and a large excess of hydrogen, offers a number of technological advantages in the pattern of separation of the cracked gas mixture.

It is known to carry out the hydrogenation process at temperatures of from 50 to 200"C, pressures of from 1 to 40 kp/cm2, preferably from 25 to 35 kp/cm2, and charges of up to 10,000 v/vh, the reactor temperature, preferably, being limited to a maximum of about 150"C.

Catalysts having palladium concentrations of up to about 0.5% by mass are used for these processes. It is known that temperature control of processes of this kind is critical in the presence of a large excess of hydrogen.

In the known technical methods of solving the problem, therefore, the temperature rise in the catalyst bed is restricted to less than 30"C, even in the presence of carbon monoxide, for example to values of under 21"C. The carbon monoxide concentration required ranges from 0.1 to about 1.0%.

In order to increase the operating stability of processes of this kind, it is known to apply a promoter to the catalysts for an increase in selectivity, e.g. silver and iron oxide. The gain in selectivity, obtained by application of a promoter, is limited and results in loss of activity.

The known processes also relate to macroporous palladium catalysts, having a low specific surface. The activity of the proposed catalysts, having surfaces of under 15 m2/g, particularly under 10 m2/g, is relatively low. Thus, a palladium catalyst, treated with silver and iron oxide promoter, on a corundum carrier, requires working temperatures of at least 100 C even in the absence of carbon monoxide. It is known that palladium/alpha - aluminium carrier catalysts with 0.045% by mass of palladium and a specific surface of 11 m2/g or less decompose only 90% of the acetylene at temperatures of from 90 to 950C, in spite of a large excess of hydrogen and absence of carbon monoxide, but they also hydrogenate from 0.5 to 1.0% of ethylene.

Furthermore, it is known to carry out the selective hydrogenation process in a tubular reactor filled with catalysts, the gas stream being passed upwards and the level of cooling agent in the reactor jacket being regulated so that, according to the particular operating conditions and properties of the catalyst, a more or less large upper portion of the catalyst bed in the tubes, optionally also in the upper vapour zone of the reactor, is not surrounded by the cooling liquid on the side of the jacket. In this way, a lower, very largely isothermally operating stage and an upper, essentially adiabatically working, stage are combined in one and the same reactor.

A considerable disadvantage of these operating principles is that special reactors or special catalysts are required, which, in comparison with conventional types, cause considerably higher costs.

Moreover, it is known that putting catalysts into operation, especially fresh ones, presents difficulties. They have a marked tendency to overheating, as a result of high olefin hydrogenation. This disadvantage can be reduced by pre-treatment of the catalyst with carbon monoxide in a special starting procedure. It is also known to improve the initial selectivity of palladium catalysts by a special hydrogen treatment at a pressure of at least 5 atmospheres and at a temperature of at least 300C above the normal working temperature.

The object of the invention consists in developing a process for the selective hydrogenation of C2-minus fractions, by which ethylene loss is reduced compared to known processes. It is a further object of the invention to widen the stability range of the hydrogenation process and to increase considerably the operating saftey of the process, compared with known processes and to shorten the starting phase and to reduce instability.

The invention seeks to devise a process for the selective hydrogenation of C2minus fractions with optimum process control, using an active palladium/carrier catalyst, having high characteristic selectively, since the main causes of ethylene losses and of the limited stability range of the process regarding the operating parameters are to be found both in the catalyst and in the conditions of the process.

It is known that palladium catalysts catalyse acetylene hydrogenation as well as ethylene hydrogenation. The high selectively of these catalysts, obtained in suitable conditions, is due to the fact that the chemisorption equilibrium, even in the presence of a large excess of olefin, lies, almost completely on the side of the acetylene, so that the latter is hydrogenated with high selectivity. If, however, the acetylene content falls below a certain limiting concentration, the absolute value of which depends, inter alia, also on the catalyst used and on the process conditions, ethylene hydrogenation sets in to an increasing extent.

As almost complete removal of acetylene is demanded, this state is reached and exceeded in any event under industrial conditions. The loss of ethylene is also increased if the course of the reaction is determined by diffusion processes on or in the catalyst blocks, in whole or in part.

Carbon monoxide increases the selectivity of the process, since it inhibits ethylene hydrogenation to a considerably greater extent than acetylene hydrogenation. This action depends on temperature. It rapidly falls with rising temperature, so that large temperature gradients in the catalyst bed cause deterioration of the selectivity and of the stability of the process and even, in the extreme case, are the cause of very high ethylene hydrogenation as a result of which the hydrogenation process passes out of control.

A further object of the invention, therefore, is to use a catalyst, which, on account of its pore structure and the distribution of the active components, very largely excludes diffusion processes and, for this reason, tolerates a large temperature gradient, in comparison with known catalysts, as well as to device the process so that the permissible temperature gradient on the catalyst side, which inevitably appears as a result of acetylene hydrogenation, is used in the process to optimum extent and, therefore, high efficiency and stability of the hydrogenation process are attained. A further object of the invention consists in devising the starting conditions so that the decomposition of acetylene required and high selectivity are attained within a short time, with process-conditioned or minimal increase in the carbon monoxide concentration.

According to this invention there is provided a process for selective hydrogenation of C2-minus fractions containing ethylene, ethane and methane and from 0.3 to 2.0% by mass of acetylene at least 1.0% by mass of hydrogen and from 1,000 to 2,000 ppm by volume of carbon monoxide, at a pressure of from I to 40 kp/cm2, a space velocity of from 2,500 to 10,000 v/vh and a reactor temperature of from 35 to 1500C, in which process an essentially macroporous palladium/carrier catalyst, optionally treated with a promoter of copper, nickel, silver and/or iron or compounds of these elements, in the form of blocks having a maximum specific surface of 50 m2/g, at least 33% of the pore volume of said catalyst being distributed among pores with radii of above 500 A.U. and having a palladium concentration from 0.005 to 0.1% by mass, is utilized within a jacket-cooled tubular reactor wherein the flow velocity of the C2-minus fraction through the catalyst tubes lies from 0.3 to 3.0 m/sec, the ratio of the internal width of the reactor tubes to the diameter or to the height respectively of the catalyst blocks lies between 8 and 25, the temperature difference between the reactor exit and reactor entry being from +600C to --150C and between any point in the catalyst bed and the cooling medium being from 0 C to +600 C, activation or re-activation of the catalyst-filled reactor being effected with a carbon monoxide concentration in the C2-minus fraction of between the process-conditioned content and 5,000 ppm by volume and at a space velocity of at least 3,000 v/vh, the reactor entry temperature being increased within from 1 to 6 hours to just below or equal to the initiating temperature of the catalyst and the required acetylene decomposition then being adjusted, optionally with an interposed interval of from 30 to 60 minutes, at a largely constant entry temperature by stepwise increases in the pressure of the cooling medium and temperature alternations during the process in the reactor being effected up to a maximum rate of 20 C/hour.

A suitable apparatus for performing the process according to the invention comprises a heat-exchanger, a pre-heater, a tubular reactor, a condenser and a green oil separator. This allows for feeding of cold raw fraction between preeater and reactor and circulation of the hot refined C2-minus fraction around the heat-exchanger. Favourable dimensions for the catalyst tubes in the reactor have proved to be an internal width of from 25 to 55 mm, preferably from 40 to 51 mm, and a length of from 4000 to 8000 mm. The cold C2-minus raw fraction is preheated in the heat-exchanger by the hot hydrogenated product stream from the reactor. The reactor entry temperature required is adjusted with the aid of the preheater and, optionally, by feeding in a partial stream of the cold raw fraction. The tubes of the reactor are filled with catalyst to about 90% by volume and are cooled from the jacket by a boiling liquid. The vapours of cooling medium are condensed in a vapour condenser and flow back into the cooling-jacket of the reactor by way of a reservoir. The level of cooling medium is regulated so that the catalyst bed in the tubes is surrounded with cooling liquid from the jacket to the full height. The reactor temperature is controlled, above all, by way of the pressure of cooling medium. The cooling medium should possess as low a boiling point as possible and a high heat of vaporisation. Cooling with liquid low-molecular hydrocarbon fractions or methanol and, at temperatures substantially above 100"C, with methanol/water mixtures has proved favourable. As a rule, the hot refined C2minus fraction first transfers part of its heat to the cold raw fraction and, after intensive cooling, the very low quantities of oligomer are removed from it in the green oil separator.

Preferably, the process according to the invention is carried out at a pressure of from 25 to 35 kp/cm2 and a space velocity of from 3000 to 7000 v/vh. The catalyst used preferably contains from 0.005 to 0.1% by mass of palladium and, in all, from 0.005 to 0.25% by mass of promoter comprising copper, nickle, silver and/or iron or their oxides on a carrier, essentially consisting of alpha-aluminium oxide, having a specific surface of from 1 to 25 m2/g, in which at least 50% of the total pore volume are distributed among pores with radii of above 500 A.U. and at most 10% on these with radii of under 100 A.U. Preferably, a catalyst is used, which possesses a specific surface of from 3 to 15 m2/g and particularly from 3 to 10 m3/g, which does not contain any pores with radii of under 100 A.U., in which at least 75% of the pore volume are distributed among pores with radii of above 500 A.U. and which contains from 0.01 to 0.05% by mass of palladium as well as from 0.003 to 0.05% by mass of nickel, copper and/or silver or their oxides and from 0.03 to 0.2% by mass of iron (III) oxide or iron (II,III) oxide. The palladium in the preferred catalyst according to the invention is located in a peripheral layer of the catalyst block, having a thickness of 1.5 mm maximum, preferably, of up to 1.0 mm, and is present, independently of the depth of penetration of the active components in the carrier block, in a relatively uniform and, under the reaction conditions, stable structure.

The preferred catalyst used contains from 0.005 to 0.05% by mass of palladium on a carrier, essentially consisting of alpha - aluminium oxide, having a specific surface of from 15 to 50 m2/g, at least 33% of the pore volume being distributed among pores with radii of above 500 A.U.

The optimum concentration of palladium in the catalyst depends on the process conditions. At high charge and high acetylene concentrations, it lies, preferably, in the upper part of the range indicated. The required operating temperature of the catalyst can be lowered through an increase in the palladium content. If the catalyst, used according to the invention, does not contain any promoters, if the palladium that it contains is distributed over the whole carrier block and/or if the specific surface is above 20 m2/g, palladium concentrations of under 0.040 by mass are preferred. With rising palladium content, rising specific surface, especially above 20 m2/g, and increasing depth of- penetration of the palladium, the ethylene balance of the process deteriorates under comparable conditions and the requirements regarding maintenance of temperature conditions increase, especially at low, usually process-conditions, carbon monoxide concentration; i.e. the stability of the hydrogenation process falls off. On application of catalysts having specific surfaces of above 50 m2/g or palladium concentrations of above 0.1% by mass, ethylene losses resulting from the hydrogenation process are unavoidable under industrial conditions, temperature control is critical and the starting behaviour is unsatisfactory. The hydrogenation process can be controlled with catalysts of this kind only with difficulty or not at all, even at very high carbon monoxide concentrations. Additionally, the selectivity of the catalyst is unfavourably influenced by micropores.

The measurement of the catalyst block depends, above all, on the internal width of the catalyst tubes. It has been found that the most favourable ratio of internal diameter of the tubes to the diameter or to the height of the catalyst blocks lies between 8 and 25, preferably between 10 and 15.

In the process according to the invention, the feed-stock flows through the catalyst tubes at a velocity of from 0.3 to 3.0 m/sec, preferably from 0.5 to 2.0 m/sec. The most important parameters for the dimensioning of the reactor tubes are the acetylene concentration in the raw fraction, the charge range to be expected and the efficiency of the cooling system. With decreasing charge of the catalyst, decreasing cooling efficiency and increasing acetylene concentration, the velocity of flow has to be raised to obtain optimum results. At acetylene concentrations of above 1% by mass, it should not fall below 0.8 m/sec and should preferably lie between 1.0 and 2.0 m/sec. Moreover, it is advantageous in these cases for sufficiently rapid removal of the heat of hydrogenation to be provided for by a relatively small tube diameter.

Best results are obtained in the process according to the invention in attempting to have as large a temperature difference as possible between reactor exit and reactor entry in the axial direction of the catalyst tubes. It has been found that this difference can be as high as +60"C with the use of the catalyst described above, without ethylene losses occurring or the hydrogenation process becoming unstable. The result is not, or not significantly, impaired even at negative values of this temperature difference of down to-I 50C, if the acetylene concentration in the C2-minus raw fraction lies below 1.0% by mass. Ethylene losses occur and the stability of the process falls off above -150C at constant carbon monoxide concentration with increasing negative temperature gradient and rising acetylene concentration, especially above 1.0% by mass.

Preferably, the entry temperature lies between 50 and 5"C below the exit temperature in the process according to the invention. The reactor entry temperature may be kept up to 200C below the initiating temperature of the catalyst, the initiating temperature being the temperature, at which the catalyst just starts to hydrogenate the acetylene under the given process conditions. During the starting phase and with catalysts of relatively low effective activity level the reactor entry temperature is preferably kept at or above the initiating temperature. In continuous operation, especially in the presence of a catalyst of high effective activity level, the entry temperature preferably lies up to 200C below the initiating temperature. The axial and radial temperature gradient between any point of the catalyst bed and the temperature of the cooling medium must not exceed +60"C.

The most favourable temperature difference for a particular case of application depends, above all, on the effective activity level of the catalyst used and on the dimensions of the reactor. For large internal widths of the catalyst tubes and with very active catalysts, it should preferably not exceed 45"C and, while fresh catalysts are put into operation, at carbon monoxide concentrations of ~2000 ppm by volume, should preferably not exceed 30"C, in the interests of ethylene yield and of process stability.

A distinctly better selectivity and considerably higher stability of the hydrogenation process are achieved by this temperature control, compared with all the prior processes, in spite of the large temperature gradients permissible.

Any alteration in the reactor entry temperature takes place in small steps, 20"C/hour, and preferably 10 C/hour, not being exceeded in all, regardless of whether raising or lowering of the temperature. The temperature of the cooling medium is similarly altered only in small steps, the alteration not exceeding, in all, 10"C/hour, and preferably 5 C/hour, in this case. With increasing operating time of the catalyst, the speed of temperature alterations can be raised.

The increase in temperature has to be effected particularly carefully and slowly, when fresh catalysts are put into operation or, within the range around and above the initiating temperature of the catalyst used, during the reactivation i.e. re starting of reactors. Furthermore, it is essential, in these cases that alterations in the temperature or in the carbon monoxide concentration are carried out under as careful analytical control of the hydrogenation process as possible. Activation, i.e.

start-up of a reactor, filled with fresh catalyst, is effected at a carbon monoxide concentration in the raw fraction, which lies between the process-conditioned value and a maximum value of 5000 ppm by volume, preferably between the process-conditioned value and 3500 ppm by volume. In the case of re-use of a catalyst, preferably, a carbon monoxide content of between the processconditioned concentration and a maximum of 3000 ppm by volume is applied. The most favourable carbon monoxide concentration for a given case depends, above all, on the dimensions of the reactor, the composition of the raw fraction, the efficiency of the cooling system as well as on the effective activity and selectivity of the catalyst used. With relatively large internal diameters of the catalyst tubes, with high acetylene concentration in the C2-minus raw fraction, with low efficiency of the cooling system, with low charge and/or with a catalyst of high activity, it is advantageous, in the interests of a short starting-time, of process stability and of ethylene yield, temporarily to increase the carbon monoxide content in the raw fraction during the starting phase, within the limits according to the invention.

fhe space velocity must not lie below 3000 v/vh during activation and, preferably, ranges from 3500 to 4000 v/vh in this phase. Advantageously, at least part of the C2-minus fraction, passing through the reactor, is recycled in the starting phase, so as to attain this charge with minimum consumption of raw fraction.

In the process, the C2-minus raw fraction is supplied to the reactor, on activation at a temperature of from 5 to 300C, preferable from 15 to 300C. Then, the entry temperature is increased in small steps to the initiating temperature within from I to 6 hours, preferably from 2 to 3 hours. Advantageously, the increase in temperature is interrupted just below or at the initiating temperature for from 30 to 60 minutes, until the temperature gradient in the reactor has stabilised.

As soon as the acetylene is being hydrogenated to a slight extent, the reactor temperature is increased in small steps by way of the pressure of cooling agent and under careful analytical control, until the acetylene content lies below 3 ppm.

After the hydrogenation process has stabilised, the carbon monoxide concentration is stepwise reduced, if this is necessary, to the process-conditioned value according to the measurement of the ethylene balance. As soon as ethylene hydrogenation becomes noticeable, the reactor temperature is lowered, parallel to the carbon monoxide content.

The process according to the invention, by comparison to the prior processes, has a number of important advantages, which have an extraordinarily favourable effect on the economy and the stability of the hydrogenation process as well as on the investment expenditure required for putting it into practice. As a result of the high characteristic selectivity of the catalysts used and of the process conditions, adjusted to the latter, a very high temperature difference is permissible in the catalyst bed or between catalyst bed and cooling medium, compared with the known processes, without the reactor tending to pass out of control or towards ethylene hydrogenation, respectively. On the contrary, the ethylene losses in the process lie distinctly below those hitherto customary. As a result, the concentration of moderators, required in the raw fraction, is low, without significant ethylene losses occurring or acetylene not being adequately hydrogenated in the case of process-conditioned deviations in the composition of the C2-minus fraction and/or in the case of variations in the process parameters. As has already been explained, this temperature difference has to be limited in the prior processes, however, to less than 30"C, in most cases even to 210C and less, in order to keep the ethylene losses within limits and to obtain an adequate stability of the process. An increased concentration of moderator is also generally required. A further substantially advantage of the process according to the invention, which results from the high temperature gradients permissible within the catalyst tubes in radial and axial direction, the high characteristic selectivity of the catalysts used and the optimum adjustment of the process conditions to the latter, lies in that it can be carried out in suitably dimensioned conventional tubular reactors and, optionally, also in intercooled bed-reactors.

Furthermore, it is advantageous for the process according to the invention to supply over a short period during the start-up a selectively hydrogenated C2-minus fraction, complying with the purity requirements, (acetylene: 3 ppm), with minimum loss of ethylene or supply of raw material, respectively, and minimum additional requirement of carbon monoxide.

Embodiments of the process according to the invention are described in the following examples: EXAMPLE I 7.2 litres of catalyst are incorporated in an experimental reactor. The reactor is part of an experimental unit, consisting of pre-heater, reactor, cooling-system and the corresponding measuring and regulating devices, including instruments for determination of parameters. The tubular reactor is surrounded by a cooling-jacket and has an internal width of 51 mm, (external diameter: 57 mm, wall thickness: 2.9 mm). The filling-level of catalyst is 3500 mm. The pre-heater is operated with low pressure steam. The cooling-system consists of the cooling-jacket of the reactor, a water-condenser and a reservoir for the cooling-liquid. Boiling methanol serves as cooling-medium.The C2-minus fraction required is taken over direct from an industrial plant for the production of ethylene. Increased carbon monoxide concentrations can be adjusted to a pre-determined value by additional feeding of carbon monoxide at the entry of the experimental unit. The reactor entry temperature is regulated by way of the pre-heater and the cooling agent temperature by way of the methanol pressure in the cooling-system by the quantity of cooling-water in the methanol-condenser. Analyses are carried out with the aid of customary gas-chromatographic methods.

The C2-minus fraction used had the following average composition in % by mass: acetylene 0.75% ethylene 50.0 ethane 13.5 hydrogen 1.65 methane 34.8 carbon monoxide 1,000 to 1,800 ppm (vol.) The calalyst is described by the following data:- apparent density (loose) 0.90 kg/litre bursting pressure 173+40 kp/cm2 composition Pd 0.034% by mass CuO 0.020% by mass Fe2O3 0.028% by mass carrier material balance structure of carrier alpha-aluminium oxide pore volume, total 0.37 cm3/g r < 100A.U. 0.0 cm3/g 100 < r < 500 A.U. 0.04 cm3/g r > 500A.U. 0.33 cm3/g specific surface 8.4 m2/g form of catalyst tablets having a diameter of 4.56 0.15 IS mm and a height of 4.6 +0.2 mm distribution of palladium ca. 0.8 mm thick peripheral layer Before the start-up, the reactor, filled with catalyst, was rendered inert with nitrogen and the carbon monoxide concentration in the C2-minus raw fraction was increased to 3000 ppm by feeding additional carbon monoxide. Then, the C2-minus fraction was supplied to the reactor at an entry temperature of 30"C and a pressure of 28.7 atmospheres. The flow velocity through the catalyst tube was 0.45 m/sec and the charge 3500 v/vh. The entry temperature was increased in small steps within 2 hours up to 600 C. The hydrogenation reaction set in at this temperature, as was proved by the rising pressure of cooling agent and the analytical results. Under analytical control and in small steps, the entry temperature was increased up to 70"C and the pressure of cooling agent to from 0.2 to 0.25 kp/cm2. Under these conditions, the acetylene content was already reduced to 2 ppm. Losses of ethylene did not occur in this case. After stable working conditions had been established, the carbon monoxide concentration was reduced stepwise to the process-conditioned value of from 1600 to 1800 ppm. In this connection, as soon as the analyses indicated small ethylene losses, the reactor temperature was also lowered in parallel.

Typical results of this phase are summarised in Table 1.

TABLE 1 C2 volume. Conditions were purposely made more stringent by a relatively high working temperature for testing the activity/time ratio of the catalyst. The process conditioned carbon monoxide concentration varied between 1000 and 1300 ppm by volume. The stability of the hydrogenation process and the tendency of the catalyst to pass out of control were tested several times, after different operating periods in analogous fashion to Example 1. Furthermore, the reactor was systematically stopped and started six times in the course of the experiment. The start-up was effected in accordance with Example 1, at charges of from 2200 to 7000 v/vh and carbon monoxide concentrations of from 1100 to 2500 ppm by volume. Typical results are summarised in Table 2. The acetylene content was < 2 ppm in all cases.

The experiment was discontinued after an effective working period of 774 hours, without a decrease in activity being recognizable by comparison with the 100th working hour.

TABLE 2 C2H4-balance, working temperature pressure of CO related to time reactor cooling agent charge concn. C2H4 used in hours entry "C kp/cm2 v/vh ppm(vol) + % 103 95 1.6 5800 2300 + 1.4 138 94 3.35 5800 2000 +0 190 95 1.15 5800 1000 +1.02 269 95 1.9 5800 2000 +1.14 404 113 2.2 5800 2500 +0.65 412 93 3.5 5800 1200 -0.55 570 98 1.5 7000 1100 +1.24 613 85 0.85 2900 1200 +0.86 647 85 4.0 2900 1000 -1.79 708 103 4.0 2200 1000 -3.18 729 95 1.5 5800 1000 +0.96 753 100 1.75 7000 1000 +0.63 770 106 2.3 5800 2500 +0.08 TABLE 3 Catalyst for Example 3 Example 4 Example 5 apparent density (loose) in kg/l 0.84 0.86 1.05 bursting pressure in kp/cm2 382+80 391+85 183+105 composition in % by mass Pd 0.032 0.025 0.035 CuO 0.018 0.017 - Fe2O3 0.044 0.06 0.1 carrier material remainder remainder remainder X-ray structure of alpha-Al2O3 alpha-Al2O3 alpha-Al2O3 carrier material with proportion of theta- and kappa-Al2O3 pore volume in cm3/g total 0.44 0.52 0.35 radii < 100 A.U. 0.0 0.03 0.07 radii between 100 and 500 A.U. 0.13 0.09 0.15 radii > 500 A.U. 0.33 0.40 0.13 specific surface in m2/g 12.0 11.0 32 TABLE 3 (cont.).

Catalyst for Example 3 Example 4 Example 5 uniform distribution peripheral peripheral throughout active components depth of depth of the whole of in the carrier penetration penetration the carrier blocks ca. 0.6 mm ca. 0.8 mm block form of catalyst tablets tablets tablets height in mm 4.70+0.15 4.75+0.15 3.4+0.1 diameter in mm 4.8 i0.2 4.75+0.15 3.6+0.2 EXAMPLE 3 The catalyst, described in Table 3, column 1, was used in the experimental unit, described in Example 1, for the selective hydrogenation of a technical C2minus fraction. In this case, the carbon monoxide concentration was varied between 1730 and 6300 ppm by volume, the charge between 2800 and 7000 v/vh, the reactor entry temperature between 67 and 104"C and the pressure of cooling agent between 0.3 and 2.55 kp/cm2. The reactor pressure varied between 26.8 and 28.3 kp/cm2.

The experiment was carried out in a way similar to that described in Example 1. The typical results are summarised in Table 4. The residual acetylene content was 2 ppm in all cases.

TABLE 4 C2H4-balance, Temperature pressure of CO-con- related to reactor entry cooling agent charge centration C2H4 used "C kp/cm2 v/vh ppm(vol) + % 86 0.55 5800 6300 +0.08 76 0.50 5800 3600 +1.22 72 0.53 5800 2500 +0.61 83 0.85 7000 2500 -2.02 80 0.94 4400 2400 -2.55 80 0.95 2900 5000 +1.05 95 2.5 5800 2600 -5.46 EXAMPLE 4 The catalyst, described in Table 3, column 2, was used in the experimental unit, described in the example 1, for the selective hydrogenation of a technical C2minus fraction. The feed-mixture contained from 1200 to 2500 ppm carbon monoxide by volume. The remaining process parameters were varied within the limits indicated in the example 3. The start-up of the reactor was effected in accordance with Example 2. The typical experimental results are summarised in Table 5; the residual acetylene concentration here was 2 ppm in all cases.

TABLE 5 C2H4-balance, Temperature pressure of CO- related to reactor cooling agent charge concn. C2H4 used "C kp/cm2 v/vh ppm(vol) + % 90 1.7 5800 1900 +1.24 90 1.6 5800 1700 +1.5 90 1.5 5800 1500 +0.83 90 1.3 4400 2100 +1.19 87 1.2 2900 2100 +1.19 93 1.55 7000 1500 +1.32 94 3.5 5800 1400 -0.24 EXAMPLE 5 The catalyst, described in Table 3, column, 3, was incorporated into the experimental reactor in accordance with Example 1. The start-up was effected in analogous fashion to Example 1, but with a carbon monoxide concentration of 1600 ppm by volume. The further execution of the experiment correspond to Example 3 and 4. The results, summarised in Table 6, are typical of this experiment. The acetylene was diminished to 3 ppm in each case. The hydrogenation reaction sets in at 57"C.

TABLE 6 C2H4-balance, temperature pressure of CO- related to reactor cooling agent charge concn. C2H4 used entry OC kp/cm2 v/vh ppm(vol) I % 72 0.35 3500 1600 +0.62 76 0.80 5800 1800 +0.94 85 0.95 5800 1200 +0.78 86 1.5 5100 1200 -0.52 85 2.5 5100 1250 -1.5 87 2.95 5800 1300 -4.36 EXAMPLE 6 Comparative Example For comparison with the results obtained in the examples, according to the invention, 1 to 5, 7.2 litres of an industrially produced catalyst that had been proved in practice for the selective hydrogenation of C2-minus fraction, were incorporated into the experimental unit, described in the Example I, and tested in accordance with the principle, used in the Examples 1 to 5, and essentially under analogous conditions. The comparative catalyst is described by the following data:- apparent density (loose) 0.75 kg/l bursting pressure 356j290 kp/cm2 composition Pd 0.034 % by mass CuO 0.01 Fe2O3 0.54 NiO 0.015 structure of carrier mixture of alpha- and theta-aluminium oxide pore volume, total 0.60 cm3/g for r < I00 AU. 0.25 cm3/g for 100 < r < 500 A.U. 0.25 cm3/g for r > 500 A.U. 0.10 cm3/g specific surface 82 m2/g form of catalyst tablets having a diameter of6.0+0.8 mm and a height of 3.2f1.1 mm On starting up, the raw fraction was placed into the reactor with a carbon monoxide content of 3800 ppm and a temperature of 25"C. The charge was 5800 v/vh. The entry temperature was increased, in small steps, up to 600C within 3 hours and kept there for I hour, until the reactor exit temperature had adjusted itself very largely. The temperature was then further increased at a speed of 6"C/hour. The hydrogenation reaction set in at 720 C. The reduction of acetylene content required, to 3 ppm, was reached about 7.5 hours after charging the reactor with the C2-minus fraction, at an entry temperature of 80"C and a pressure of cooling agent of 0.8 kp/cm2. The pressure in the reaction varied between 26 and 28 kp/cm2. The remaining parameters were varied within the limits indicated in the example 1. The typical values, summarized in Table 7, were obtained.

TABLE 7 C2H4-balance, temperature pressure of CO- related to reactor cooling agent charge concn. C2H4 used entry "C kp/cm2 v/vh ppm(vol) + % 82 0.8 5800 1800 +0.20 80 0.8 5800 2220 -0.32 75 0.8 2900 1600 -0.75 82 1.5 5800 1500 -3.66 82 2.0 5800 1500 -41.0 (complete consumption of H2) 74 1.2 2800 1500 -4.29 105 0.7 2800 1600 -1.58 84 0.8 7000 1500 -1.5 Start-up of the catalyst with carbon monoxide concentrations of less than 2000 ppm was not possible, as it had a strong tendency to pass out of control. Besides the process stability attainable, compared with the Examples 1 to 5, was considerably lower. Even a relatively small increase in the reactor temperature, especially by way of the pressure of cooling agent, of a lowering of the carbon monoxide concentration led to rapidly rising losses of ethylene and, finally, even to loss of control of the reactor. This becomes particularly evident on comparing the results, summarised in Table 8, regarding process stability and loss of control of the catalysts.

TABLE 8 C2H4-balance, Temperature Pressure of Co- related to reactor cooling agent Charge concn. C2H4 used Example entry OC kp/cm2 v/vh ppm(vol) I % 70 4.0 3500 1700 +0.0 45 3.95 5800 1800 +1.63 2 93 3.50 5800 1200 -0.55 85 4.0 2900 1000 -1.79 103 4.0 2200 1000 -3.18 3 95 2.5 5800 2600 -5.46 4 94 3.5 5800 1400 -0.24 5 87 2.95 5800 1300 -4.36 6 82 1.5 5800 1500 -3.66 82 2.0 5800 1500 -41.0 (complete hydrogen ation) WHAT WE CLAIM IS: 1. A process for selective hydrogenation of C2-minus fractions containing ethylene, ethane and methane and from 0.3 to 2.0% by mass of acetylene, at least 1.0"it by mass of hydrogen and from 1,000 to 2,000 ppm by volume of carbon monoxide, at a pressure of from 1 to 40 kp/cm2, a space velocity of from 2,500 to 10,000 v/vh and a reactor temperature of from 35 to 150"C, in which process an essentially macroporous palladium/carrier catalyst, optionally treated with a promoter of copper, nickel, silver and/or iron or compounds of these elements, in the form of blocks having a maximum specific surface of 50 m2/g, at least 33 /n of the pore volume of said catalyst being distributed

**WARNING** end of DESC field may overlap start of CLMS **.

Claims

**WARNING** start of CLMS field may overlap end of DESC **.

TABLE 7 C2H4-balance, temperature pressure of CO- related to reactor cooling agent charge concn. C2H4 used entry "C kp/cm2 v/vh ppm(vol) + %

82 0.8 5800 1800 +0.20

80 0.8 5800 2220 -0.32

75 0.8 2900 1600 -0.75

82 1.5 5800 1500 -3.66

82 2.0 5800 1500 -41.0 (complete consumption of H2)

74 1.2 2800 1500 -4.29

105 0.7 2800 1600 -1.58

84 0.8 7000 1500 -1.5 Start-up of the catalyst with carbon monoxide concentrations of less than 2000 ppm was not possible, as it had a strong tendency to pass out of control. Besides the process stability attainable, compared with the Examples 1 to 5, was considerably lower. Even a relatively small increase in the reactor temperature, especially by way of the pressure of cooling agent, of a lowering of the carbon monoxide concentration led to rapidly rising losses of ethylene and, finally, even to loss of control of the reactor. This becomes particularly evident on comparing the results, summarised in Table 8, regarding process stability and loss of control of the catalysts.

TABLE 8 C2H4-balance, Temperature Pressure of Co- related to reactor cooling agent Charge concn. C2H4 used Example entry OC kp/cm2 v/vh ppm(vol) I %

70 4.0 3500 1700 +0.0

45 3.95 5800 1800 +1.63

2 93 3.50 5800 1200 -0.55

85 4.0 2900 1000 -1.79

103 4.0 2200 1000 -3.18

3 95 2.5 5800 2600 -5.46

4 94 3.5 5800 1400 -0.24

5 87 2.95 5800 1300 -4.36

6 82 1.5 5800 1500 -3.66

82 2.0 5800 1500 -41.0 (complete hydrogen ation) WHAT WE CLAIM IS: 1. A process for selective hydrogenation of C2-minus fractions containing ethylene, ethane and methane and from 0.3 to 2.0% by mass of acetylene, at least 1.0"it by mass of hydrogen and from 1,000 to 2,000 ppm by volume of carbon monoxide, at a pressure of from 1 to 40 kp/cm2, a space velocity of from 2,500 to 10,000 v/vh and a reactor temperature of from 35 to 150"C, in which process an essentially macroporous palladium/carrier catalyst, optionally treated with a promoter of copper, nickel, silver and/or iron or compounds of these elements, in the form of blocks having a maximum specific surface of 50 m2/g, at least 33 /n of the pore volume of said catalyst being distributed

among the pores with radii of above 500 A.U. and having a palladium concentration from 0.005 to 0.1% by mass, is utilized within ajacket-cooled tubular reactor wherein the flow velocity of the C2-minus fraction through the catalyst tubes lies from 0.3 to 3.0 m/sec, the ratio of the internal width of the reactor tubes to the diameter or to the height respectively of the catalyst blocks lies between 8 and 25, the temperature difference between the reactor exit and reactor entry being from +600C to -15"C and between any point in the catalyst bed and the cooling medium being from OOC to +600 C, activation or re-activation of the catalyst-filled reactor being effected with a carbon monoxide concentration in the C2-minus fraction of between the process-conditioned content and 5,000 ppm by volume and at a space velocity of at least 3,000 v/vh, the reactor entry temperature being increased within from I to 6 hours to just below or equal to the initiating temperature of the catalyst and the required acetylene decomposition then being adjusted, optionally with an interposed interval of from 30 to 60 minutes, at a largely constant entry temperature by stepwise increases in the pressure of the cooling medium and temperature alterations during the process in the reactor being effected up to a maximum rate of 20"C/hour.
2. Process according to Claim I, wherein the catalyst contains in total from 0.0005 to 0.25% by mass of copper, nickel, silver and/or iron or their oxides on a carrier, essentially consisting of alpha-aluminium oxide, having a specific surface of from I to 25 m2/g and in which at least 50% of the total pore volume are distributed among pores with radii of above 500 A.U. and a maximum of 10% on those with radii of under 100 A.U.
3. Process according to Claim I or 2, wherein the catalyst used has a specific surface of from 3 to 15 m2/g, and preferably from 3 to 10 m2/g, does not possess any pores with radii of under 100 A.U. and at least 75% of the pore volume are distributed among pores with radii of above 500 A.U.
4. Process according to Claim 1 or 2, wherein the catalyst used contains from 0.01 to 0.05% by mass of palladium and, as promoters, from 0.003 to 0.05 zÓ by mass of nickel, copper and/or silver or their oxides and from 0.03 to 0.2% by mass of iron (III) oxide or iron (II,III) oxide.
5. Process according to any of Claims 1, 2 or 4, wherein the catalyst used contains the palladium in a peripheral layer of the catalyst block, having a maximum thickness of 1.5 mm and, preferably having a maximum thickness of 1.0 mm.
6. Process according to any of Claims 1, 2 or 5, wherein the catalyst used contains from 0.005 to 0.5% by mass of palladium on a carrier, essentially consisting of alpha-aluminium oxide, having a specific surface of from 15 to 50 m2/g, at least 33% of the pore volume being distributed among pores with radii of above 500 A.U.
7. Process according to Claim 1, wherein the ratio of internal diameter of a catalyst tube to the diameter or height, respectively, of the catalyst blocks lies between 10 and 15.
8. Process according to Claim 1, wherein the flow velocity of the C2-minus fraction through the catalyst tubes is from 0.5 to 2.0 m/sec.
9. Process according to Claim I, wherein the temperature at any point in the catalyst bed is up to 450C (maximum) above the temperature of the cooling medium.
10. Process according to Claim I to 9, wherein the difference between reactor exit temperature and reactor entry temperature lies from +50 to +50C.
11. Process according to Claim I or 10, wherein the reactor entry temperature is kept at up to 200C below the initiating temperature of the catalyst used, the initiating temperature being the temperature, at which the catalyst starts to hydrogenate the acetylene under the given process conditions.
12. Process according to any of Claims 1, 10 or 11, wherein upon introducing fresh catalysts or re-introducing catalysts, the carbon monoxide concentration in the raw fraction lies between the process-conditioned value and a maximum value of 3500 ppm by volume, the space velocity is at least 3000 v/vh, preferably from 3500 to 4000 v/vh, and the reactor entry temperature is increased in small steps within from I to 6 hours, preferably from 2 to 3 hours, to just below or to the initiating temperature of the catalyst.
13. Process according to Claim 1 or 12, wherein upon re-introducing a catalyst, the carbon monoxide concentration in the raw fraction lies between the processconditioned content and 3000 ppm by volume.
14. Process according to any of Claims 1, 10 or 11, wherein the reactor entry temperature is altered where required in small steps at a maximum total rate of 200hour and preferably at a rate of 10 C/hour.
15. Process according to Claim 1 or 9, wherein the temperature of the cooling medium is altered where required in small steps, at a maximum total rate of 10"C/hour and preferably at a maximum total rate of 5 C/hour.
16. A process for the selective hydrogenation of C2-minus fractions carried out substantially as herein described and exemplified in any of Examples 1 to 5.