CS206644B1 - Selective hydrogenation method of unsaturated c2-fractions - Google Patents

Description

(54) A method for the selective hydrogenation of unsaturated _C2 - fractions

The invention relates to a method of selectively hydrogenaoe unsaturated C ₂ -fractions, a heterogeneous,. Catalyst, a chemical process which, in the process of recovering ethylene from the gasoline pyrolysis fission gases, maximally eliminates the acetylene contained in the crude fraction and interferes with olefin processing without substantially losing ethylene.

The production of ethylene as an intermediate for the production of plastics has reached enormous technical importance. It is preferably obtained by separating and refining the fission gases from the pyrolysis of gasoline. Selective hydrogenation of the acetylenic bond has proved to be an almost universal method of refining. Such selective hydrogenation processes are required to remove the acetylene present as a rule in the crude fraction in an amount of 0.3 to 2.0% to less than 3 ppm without hydrogenating substantial amounts of ethylene present in a large excess.

For energy reasons and in order to keep the polymer formation volume on the catalyst as low as possible, generally the lowest working temperature is required because, as is well known, polymer formation, independently of the catalyst system used, becomes more and more favorable as temperature increases with catalyst deactivation. In addition, losses due to excess hydrogen with increasing temperature increase extremely rapidly, even when moderators, such as ultraviolet oxide, are present, increasing selectivity. Nevertheless a selective hydrogenation of acetylene in the so-called. -Fractions unsaturated C ₂ to paláidiových catalysts, i.e., in a batch of the product, which in addition to acetylene, ethylene and ethane, more carbon monoxide, methane and a large excess of hydrogen, a number of technological advantages in adjusting the separation of the mixture of the cracked gas .

It is known that the hydrogenation process is carried out at temperatures from 50 to 200 ° C and pressures of 1, 10 ⁵ to 40 105 _{Pa> preferably} 25.10 ⁵ and 35.10 ⁵ Pa and loads up to 10 000 V / wh, wherein the reactor temperature is delimited upwards by about 150 [deg.] C.

Catalysts with a palladium concentration of up to about 0.5% by weight are used for this process. Temperature management of such processes in the presence of a large excess of hydrogen is known to be critical (V. K. Lam and L. Lloyd, Oil Gas Journal 1972) March 27 (66-70).

For this reason, in known technical solutions, the temperature rise in the catalyst bed, even in the presence of carbon monoxide, is limited to below 30 ° C, somewhere below 21 ° C. The required concentration of carbon monoxide is between 0.1 and about 1.0% ·

In order to increase the operational stability of such processes, it is known that promoters, for example silver and iron oxides or copper and silver, are treated with catalysts to increase selectivity. The increase in selectivity achieved by promoter processing is limited and associated with a decrease in activity.

The known methods also relate to macroporous, surface-poor palladium catalysts. However, the activity of the proposed catalysts with surfaces below 15 m ² / g, in particular below 10 m ² / g, is relatively low. Thus, the palladium carrier on the corundum carrier, treated with silver and iron oxide promoters, requires itself, in the absence of carbon monoxide, an operating temperature of at least 100 ° C. It is shown in the patent literature that supported palladium-alpha aluminum catalysts with 0.0045 wt. palladium and with a specific surface area of 11 m ² / g or less, despite a large excess of hydrogen and the absence of carbon monoxide, only 90% of the acetylene is decomposed at 90 to 95 ° C, but already 0.5 to 1.0% of the ethylene is hydrogenated.

It is also known to carry out a selective hydrogenation process in a catalyst-filled tubular reactor, wherein the gas flow is conducted from the bottom up and the coolant state in the reactor jacket is controlled so that depending on the operating conditions and catalyst properties Even in the upper vat space of the reactor, the sides of the shell were not surrounded by coolant. In this way, the lower, maximum isothermally guided stage and the upper substantially adiabetically operating stage are combined in one and the same reactor. A considerable drawback of these process principles is that they are substantially more expensive than conventional types.

In addition, it is known that starting fresh catalysts, in particular, is problematic. These tend to overheat due to high olefin hydrogenation (WK Lam and L. Lloyd, Oil Gas J. 1972) March 27 (66-77). This deficiency can be alleviated by pretreating the catalyst with carbon monoxide in a special start-up process. It is also known that the initial selectivity of palladium catalysts is improved by special treatment with hydrogen at a pressure of at least ^{5 bar} and a temperature which is at least 30 ° C above normal operating temperature.

Objective of the invention consists in that it should provide a method for the selective hydrogenation of unsaturated C ₂ -fractions whereby diminishing over known methods ethylene losses. In addition, it is an object of the present invention to increase the stability of the hydrogenation process, thereby substantially increasing the process safety of the process as compared to known methods, as well as shortening the start-up phase and reducing this characteristic instability.

The invention object of the provide a process for the selective hydrogenation of unsaturated C ₂ -fractions with optimum process using a palladium supported catalyst having high intrinsic selectivities because the main cause of the loss of ethylene and reduced the stability of the process with regard to operating parameters is based both on the catalyst and both in process conditions. As is known, catalysis of palladium catalysts by both hydrogenation of acetylene and ethylene. The high selectivity of these catalysts under suitable conditions is achieved by the fact that the chemisorption equilibrium, even in the presence of a large excess of olefins, is practically entirely on the acetylene side, so that it is hydrogenated with high selectivity. However, if the acetylene content falls below a certain limit concentration, the absolute value of which, inter alia, also depends on the catalyst used and the process conditions, the hydrogenation of ethylene occurs to an increasing extent.

This state is achieved and exceeded under technical conditions in any case, since complete acetylene removal is required. The loss of ethylene also increases when diffusion processes on or in the catalyst moldings determine or co-determine the course of the reaction.

Carbon monoxide increases the selectivity of the process by preventing substantially more hydrogenation of ethylene than acetylene. Its (temperature-dependent effect, however, decreases rapidly with increasing temperature, so that large temperature gradients in the catalyst bed cause a deterioration in selectivity and process stability, and in extreme cases even cause very high ethylene hydrogenation, leaving the hydrogenation process out control.

It is an object of the present invention to provide a catalyst which, by virtue of its pore structure and the distribution of the active ingredients, avoids maximum diffusion processes and thereby permits, without reducing its activity, large temperature gradients compared to known catalysts, In order to achieve a high efficiency and stability of the hydrogenation process, the temperature gradient permissible with respect to the catalyst necessarily occurring as a result of the hydrogenation of acetylene is utilized optimally for the process. A further object of the invention is that the start-up conditions are to be arranged in such a way that the desired acetylene removal and high selectivity are achieved in a short time under process conditions or with a minimum increase in carbon monoxide concentration.

For the selective hydrogenation of an unsaturated C ₂ - fraction, "which contains 0.3 to 2.0 wt%. acetylene, at least. 1 wt. hydrogen, and generally 1,000 to 2,000 vol. ppm of carbon monoxide, in a tubular reactor with cooling jacket, at a pressure of 1 · 10 ⁵ to ⁴⁰ 10 ⁵ Pa, a space velocity of 2,500 to 10,000 v / wh and a reactor temperature of 35 to 150 ° C To solve these problems according to the invention, a catalyst-filled tubular reactor having a concentration of carbon monoxide in unsaturated C2 - (fraction between the content of the conditional process up to 5000 v / ppm and a space velocity of at least 3000 v / vh) putting into operation using macroporous and optionally 0.0005 to 0.25% by weight of copper, nickel, silver and / or iron, or a compound of these elements, a supported palladium catalyst with a specific surface area of at most 50 m ² / g, wherein at least 33% porosity with pores attributable to a radius greater than 500 AE palladium concentration from 0.005 to 0.1 wt.%, whereby the flow velocity of unsaturated C ₂ -fractions tube catalysis or is between 0.3 to 3.0 m / s and the ratio of the reactor head tube to the diameter and / or height of the catalyst moldings is between 8 to 25 and the inlet temperature of the reactor is increased to near below the catalyst start-up temperature within 1 to 6 hours; to this temperature, and optionally, with intermediate interruptions of 30 to 60 minutes, at substantially constant inlet temperature, the desired coolant removal is set by gradually increasing the pressure of the coolant and the temperature changes in the reactor are carried out at a maximum rate of 20 ° C / h. from the reactor and the inlet to the reactor as well as between any point in the catalyst bed and the cooling medium is 0 to 60 ° C.

A preferred embodiment of the process according to the invention comprises a heat exchanger, a preheater, a tubular reactor, a cooler and an oil separator. Besides allowing supply of cold raw fraction between the preheater and the reactor and the heat exchanger bypass hot refined unsaturated ° C for ₂ -fractions. A favorable dimensioning of the catalyst tube in the reactor has been shown to be a clear width of 25 to 55 mm, preferably 40 to 51 mm and a length of 4000 to 8000 mm. In the heat exchanger, the cold unsaturated C ₂ fraction is preheated by a hot stream of hydrogenated product leaving the reactor. The desired inlet temperature of the reactor is set by means of a preheater and, if necessary, the addition of a partial stream of cold raw fraction. The reactor tubes are about 90% filled with catalyst and are cooled by boiling liquid from the sides of the jacket. The refrigerant vapor is condensed in the vapor cooler and flows through the reservoir back into the reactor jacket. The state of the coolant is controlled so that the bed of catalyst in the tubes is completely surrounded by coolant from the side of the housing. The temperature of the reactor is controlled primarily by means of refrigerant pressure. The refrigerant should have a low boiling point and a high evaporation heat. Cooling with liquid, low molecular weight hydrocarbon or methanol fractions and at temperatures well above 100 ° C with methanol-water mixtures has proven to be beneficial. The hot refined unsaturated C ₂ -fractions gives normally first part of its heat to the cold raw fraction and after intensive cooling is removed from it in the oil separator Gruenově very small amounts of oligomers.

Preferably the process is conducted at a pressure of 25.10 ⁵ Pa and 35.10 ⁵ Pa and a space velocity of 3000 to 7000 v / vh. The catalyst support used consists largely or entirely of alpha alumina. The pores should not consist of more than 10% of their volume from pores with a radius below 100 AE. The catalyst contains a total of 0.0005 to 0.25 wt. promoters consisting of iron or elements of the first side group of the Mendeleyov periodic system. Preferably, a catalyst having a specific surface area of 3 to 15 m ² / g and in particular 3 to 10 m ² / g is used, which contains no pores with radii below 100 AE, in which at least 75% of the pore volume AE, which contains 0.01 to 0.05 wt. % palladium as well as 0.003 to 0.05 wt. % nickel, copper and / or silver or their oxides and 0.03 to 0.2 wt. ferric oxide and ferrous oxide. In addition, the palladium in the preferred catalyst of the present invention is present in a maximum of 1.5 mimes, but preferably up to 1.0 mm, of the peripheral layer of the catalyst compact and is present independently of the depth of penetration of the active components in the catalyst support compact. reaction stable structure.

The optimum palladium concentration in the catalyst is governed by the process conditions. At high loads and high acetylene concentrations, it preferably moves at the top of said region. In addition, the required operating temperature of the catalyst can be reduced by increasing the palladium content. If the catalyst used according to the invention contains no promoters, the palladium present therein is distributed throughout the support and / or if the specific surface area is above 20 m ² / g, palladium concentrations below 0.040% by weight are preferred. With increasing palladium content, increasing specific surface area, especially above 20 m ² / g and increasing penetration depth of palladium, the ethylene balance of the process deteriorates under comparable conditions and the requirements with respect to temperature compliance, especially at lower, normally, process that determines the concentration of carbon monoxide, ie the stability of the hydrogenation process decreases. When using catalysts with specific surfaces above 50 m ² / g or palladium concentrations above 0.1% by weight. the loss of ethylene by the hydrogenation process under technical conditions is unavoidable, the temperature conduction is critical and the start-up behavior is unsatisfactory. Hydrogenation processes with such catalysts, even at very high concentrations of carbon monoxide, are difficult to control or cannot control at all. In addition, the selectivity of the catalyst is adversely affected by micropores.

The size of the catalyst compacts is determined primarily by the clear width of the catalyst tube. It has been found that the most favorable ratio of the inner diameter of the tube to the diameter or height of the catalyst compact is between 8 and 25, preferably between 10 and 15.

According to the process of the invention, the feed product flows at a rate of 0.3 to 3.0 m / s, preferably 0.5 to 2.0 m / s, through the catalyst tubes. The most important determinants for sizing the reactor tubes are the acetylemic concentration in the crude fraction, the load area to be expected and the cooling system performance. As the catalyst load decreases, the cooling capacity decreases and the acetylene concentration increases, the flow rate must be increased for optimum results. At acetylene concentrations above 1 wt. In such cases, it is furthermore advantageous if a sufficiently rapid dissipation of the hydrogenation heat with a comparatively small pipe diameter is ensured.

The best results were achieved using the process according to the invention when a temperature difference between the reactor outlet and the reactor inlet was sought as much as possible. It has been found that using the catalyst described above, this difference can be as high as 4-60 ° C without loss of ethylene or without the hydrogenation process becoming unstable. Even at negative values of this temperature difference up to -15 ° C, the result was not affected or not significantly affected when the acetylene concentration in the unsaturated crude C ₂ fraction was below 1.0 wt%. Above -15 ° C, acetylene losses, in particular above 1.0% by weight, occurred at a constant concentration of carbon monoxide with increasing negative temperature gradients and increasing concentrations of acetylene, especially above 1.0% by weight. losses of ethylene and decreased process stability. Preferably, the inlet temperature of the process of the invention is 50 to 5 ° C below the outlet temperature. In addition, according to the invention, the inlet temperature of the reactor is kept up to 20 ° C below the catalyst start-up temperature, the catalyst start-up temperature being the temperature at which the catalyst just begins to hydrogenate acetylene under given process conditions. During the start-up phase and for catalysts with a relatively low current activity level, the reactor inlet temperature is preferably maintained at or above the start-up temperature. In continuous operation, especially in the presence of a catalyst with a high current activity level, the inlet temperature is preferably up to 20 ° C below the lowering temperature. According to the invention, the axial and radial temperature gradient between any point in the catalyst bed with the coolant temperature must not exceed + 60 ° C. The temperature difference most favorable to the particular application depends mainly on the actual activity level of the catalyst used and the reactor dimensions. At large clear tube widths for catalysts and very active catalysts, it should preferably not exceed 45 ° C for ethylene yield and process stability, and during start-up of fresh catalysts at concentrations of 2000 ppm carbon monoxide preferably 30 ° C.

In particular, this temperature control, in comparison with the known methods, despite permissible large temperature gradients, results in a markedly better selectivity and a considerably higher stability of the hydrogenation process.

Each change in the inlet temperature of the reactor is carried out in small steps, generally not exceeding 20 ° C / h and preferably 10 ° C / h, irrespective of whether the temperature is increased or decreased. The temperature of the refrigerant according to the invention also changes only in small steps, in which case the change does not exceed in total 10 ° C / h and preferably 5 ° C / h. As the catalyst operating time increases, the rate of temperature change may increase.

Particularly carefully, the temperature must be increased when starting fresh catalysts, optionally in the region around and above the starting temperature of the catalyst used when restarting the reactors. It is also essential that changes in the temperature or concentration of carbon monoxide are carried out with the greatest possible analytical control of the hydrogenation process.

According to the invention, the start-up of the fresh catalyst catalyst is carried out with a concentration of carbon monoxide in the crude fraction which varies between the conditioning process and a maximum of 5000 ppm, preferably between the conditioning process and 3500 ppm. When restarting the catalyst, the carbon monoxide content is preferably used between the conditional concentration process and a maximum of 3000 ppm by volume. The concentration of carbon monoxide that is most favorable for the particular case depends primarily on the reactor design, the composition of the crude fraction, the performance of the cooling system as well as the actual activity and selectivity of the catalyst used. At relatively internal diameters of the catalyst tubes, at high acetylene concentration in the unsaturated C ₂ - crude fraction, at low cooling system performance, at low load and / or at high activity catalyst, stability is of interest for short start-up times process and ethylene yield advantageously increase the carbon monoxide content of the crude fraction during the start-up phase transiently within the limits determined according to the invention.

The space velocity must not be below 3,000 v / vh when starting the process. Preferably, during the start-up phase, at least a portion of the unsaturated C ₂ fraction passing through the reactor is returned to achieve this load with minimum consumption to the crude fraction.

When putting into operation the process according to the invention, the unsaturated C ₂ -fractions fed to the reactor at 5 to 30 ° C, preferably 15-30 ° C. Thereafter, the inlet temperature is increased in small steps up to the lowering temperature within 1 to 6 hours, preferably within 2 to 3 hours. Preferably, the temperature rise is shortly below or at the starting temperature for 30 to 60 minutes until the temperature gradient in the reactor is as equal as possible. Once the acetylene has been hydrogenated to a small extent, the reactor temperature rises in small increments above the refrigerant pressure and under careful analytical control until the acetylene content is below 3 ppm. After the hydrogenation process has stabilized, the carbon monoxide concentration is reduced, if necessary, according to the ethylene balance data, down to the process conditioning value. As soon as the hydrogenation of ethylene becomes apparent, the reactor temperature decreases in accordance with the carbon monoxide content.

The process according to the invention has a number of significant advantages over the known processes, which have an extremely favorable effect on the economy and stability of the hydrogenation process, as well as the investment costs required for its implementation. Owing to the inherent high selectivity of the catalysts used and the conditions which are to be accepted in this manner, a very high temperature difference in the catalyst bed or between the catalyst bed and the cooling medium is permissible compared to the known processes without the reactor being prone to separate or hydrogenate ethylene. By contrast, ethylene losses in the process according to the invention are clearly below conventional losses. This moderator concentration necessary in the raw fraction lower, without causing variations in the composition of an unsaturated C ₂ -fractions, conditional process and / or variations in the process parameters to significant losses of ethylene or acetylene is sufficiently nehydrogenoval. In the known methods, on the other hand, as already mentioned, this temperature difference must be limited to a temperature difference below 30 ° C, in most cases even 21 ° C and below, in order to keep the ethylene losses within limits and achieve sufficient stability process. As a rule, an increased moderator concentration is also required.

Another important advantage of the process according to the invention, which results from the high permissible temperature gradients inside the catalyst tubes in the radial and axial direction, the high intrinsic selectivity of the catalysts used and the conditions corresponding to this process, is that it can be carried out in conventional suitably sized tubular reactors. and possibly even in-process cooled bed reactors, with the same or better efficiency with respect to operating results than with the principles of the methods disclosed in the patent literature, which require special reactors with high investment costs and complicated process management with increased failure rate. It is furthermore advantageous that the process according to the invention provides a selectively hydrogenated unsaturated C ₂ fraction of corresponding purity i (acetylene 3 ppm) after a short start-up time, with minimum losses of ethylene or feedstock and minimum additional consumption of carbon monoxide.

Several useful and preferred embodiments of the process according to the invention are described in the following examples.

Example 1

7.2 l of catalyst are charged into the test reactor. A reactor is a part of a test facility consisting of a preheater, a reactor, a cooling system and corresponding measuring and control equipment, including instruments for indicating measured quantities. The tubular reactor is surrounded by a cooling jacket and has a clear width of 51 mm (outer diameter 57 mm, wall thickness 2.9 mm). The catalyst charge height is 3500 mm. The preheater is powered by low pressure steam. The cooling system consists of a reactor cooling jacket, a water cooler and a coolant reservoir. Boiling methanol serves as a cooling medium. Required unsaturated C ₂ -fractions were taken directly from the technical equipment for the production of ethylene. Increased carbon monoxide concentrations can be deliberately adjusted by additionally supplying carbon monoxide at the inlet of the test apparatus.

The reactor inlet temperature is controlled through the preheater and the coolant temperature by the methanol pressure in the cooling system using the amount of cooling water in the methanol condenser. Analysis is carried out by conventional chromatographic methods.

Use of an unsaturated C ₂ -fractions had the following average composition in% by wt .:

acetylene 0.75% ethylene 50.0 ethane 13.5 hydrogen 1.65 methane 34.8 carbon monoxide 1 000 to 1 800 ppm / volume

The catalyst is characterized by the following data:

bulk density burst composition carrier structure porosity, total r 100 A 100r500 A r 500 A specific surface catalyst form palladium distribution

0.90 kg / 1 173 + 40.10 ⁵ Pa

Pd 0.034 wt. CuO 0.020 wt. % Fe ₂ O ₃ 0.028 wt. support material residue alpha alumina

0.37 cm ³ / g 0.0 cm ³ / g 0.04 cm ³ / g 0.33 cm ³ / g 8.4 m ² / g tablets with a diameter

4.56 and 0.15 mm and a height of 4.6 ± mm about 0.8 mm thick peripheral layer

Prior to commissioning, the catalyst-filled reactor was purged with inert nitrogen gas, and the unsaturated crude C ₂ fraction was increased by an additional addition of carbon monoxide to a concentration of 300 ppm of carbon monoxide. It was then at an inlet temperature of 30 ° C and a pressure of 28.7. 10 ⁵ Pa is introduced into the reactor unsaturated C2-fraction. The catalyst tube flow rate was 0.45 m / s and the load was 3,500 v / vh. In small steps, the inlet temperature was raised to 60 ° C over 2 hours. At this temperature, the hydrogenation reaction started, as evidenced by increasing refrigerant pressure and analysis results. Under analytical control, the inlet temperature was increased to 70 ° C and the polishing pressure to 0.2.10 ⁵ to 0.25 in small increments. 10 ⁵ Pa. Under these conditions, the acetylene was already decomposed to 2 ppm. Ethylene losses did not occur. After stable operating conditions were established, the carbon monoxide concentration was gradually reduced to a process conditioning value, i.e. 1600 to 1800 ppm. At the same time, once the analyzes showed a small loss of ethylene, the reactor temperature was lowered simultaneously.

The characteristic results of this phase are summarized in Table 1, lines 1 to 4.

Then the process conditions were changed within the following limits:

Pressure 28.10 ^s to 30.10 ⁵ Pa temperature 45 to 120 ° C refrigerant pressure 1,05.10® to 4,75.10® Pa load 2,200 to 7,000 v / vh

Table 1

temperature entry reactor Noc: 2 ° C coolant pressure Pa / cm ² load v / vh ended. WHAT vol - ppm balance C ₂ H ₄ (relative to the stake 0, ¾) ±% 1 2 3 4 5 70 0,2 .10® 3 500 3 000 (-1.15 70 0,25.10® 3 500 2 500 -0.98 70 0,25.10® 3 500 2 000 -1.27 70 0,25.10® 3 500 1700 -0.70 70 4,0 .10® 3 500 1700 ; 0.0 72 0,3 .10® 4 400 1700 -1.41 82 0,8 .10® 5 800 2 000 -0.82 82 0,75.10® 5 800 1700 - 0,71 88 1,25.10® 7 000 1700 -0.29 60 0.90.10® 5 800 1700 -0.66 50 0.90.10® 5 800 1700 -1.34 45 0,92.10® 5 800 1800 -1.28 45 2,3 .10® 5 800 1 800 -1.06 45 3,95.10® 5 800 1800 b 1.63

The process stability and the tendency of the catalyst to "run" (i.e., prohydrogenation) were tested, among other things, at a constant, process conditional on the concentration of carbon monoxide, by systematically varying the reactor temperature by means of refrigerant pressure. The characteristic values are summarized in Table 1. After an operating time of 142 hours, the process was interrupted and restarted under conditions substantially analogous to those described above. The concentration of carbon monoxide in the crude fraction was 2050 ppm. Stable operating conditions were reached after 4.5 hours. The residual acetylene content was about 1 ppm. Ethylene yields of between 0.93 and 1.4% were achieved based on the ethylene charged.

Example 2

7.2 liters of the catalyst of Example 1 were charged to the experimental apparatus described in Example 1. The start-up was performed analogously to Example 1 with a load of 5,800 v / vh, a catalyst tube flow rate of 0.8 m / s and an oxide concentration. carbon monoxide 2,500 vol.- ppm. To test the catalyst activity over time, conditions were relatively deliberately sharpened by the relatively high operating temperature. The concentration of the carbon monoxide conditioning the process varied between 1,000 to 1,300 vol / ppm. The stability of the hydrogenation process and the running of the catalyst were tested several times after various working times in analogy to Example 1. Further, the reactor was stopped and started six times in the course of the experiment. The start-up was carried out according to Example 1 at loads of 2,200 to 7,000 v / vh and a carbon monoxide concentration of 1,000 to 2,500 vol / ppm. The characteristic results are summarized in Table 2. The residual acetylene content was in all cases <2 ppm. After an effective operating time of 774, the experiment was discontinued without a noticeable loss of activity over the 100th operating hour.

Table 2

operating time in hours reactor inlet temperature ° C • + 2 u load v / vh conc. CO in vol. - ppm o ** i — Η Γ) íbď + i + λ (B) 103 95 1,6 .10® 5 800 2 300 + 1,4 138 94 3,35.10® 5 800 2 000 ± o 190 95 1,15.10® 5 800 1000 + 1.02 269 95 1.9 .10® 5 800 2 000 + 1.14 404 113 2,2 .10® 5 800 2 500 + 0.65 412 93 3.5 .10® 5 800 1200 -0.55 570 98 1,5 .10® 7 000 1100 + 1.24 613 85 0,85.10® 2 900 1200 + 0.86 647 85 4,0 .10® 2 900 1000 -1.79 708 103 4,0 .10® 2 200 1000 -3.18 729 95 1,5 .10® 5 800 1000 -0.96 753 100 ALIGN! 1,75.10® 7 000 1000 -0.63 770 106 2,3 .10® 5 800 2 500 h 0.08

Example 3

The catalyst, as described in Table 3, was used in the experimental apparatus described in Example 1 to selectively hydrogenate the technical unsaturated C2-fraction. The carbon monoxide concentrations were varied between 1,730 to 6,300 vol / ppm. load between 2800-7000 v / vh, reactor inlet temperature between 67-104 ° C and refrigerant pressure between 0.3. 10® to 2.55.10® Pa / cm ² . The pressure in the reactor varied between 26.8 x 10 3 and 28.3 x 10 3 Pa / cm ² .

The experimental design was as described in Example 1. The characteristic results are summarized in Table 4. The residual acetylene content was in all cases 2 ppm.

Table 3

catalyst for Example 3 Example 4 Example 5 bulk density in kg / l 0.84 0.86 1.05 burst in Pa / cm ² (382 ± 80). . 10® Pa (301 ± 85). . 10® Pa (185 +105). . 10® Pa % composition Pd CuO Fe ₂ O ₃ 0,032 0.018 0,044 0,025 0.017 o, o§ 0,035 0.1 material carriers residue residue residue X-ray structure carriers alpha AI2O3 alpha A1 ₂ O ₃ alpha A1 ₂ O ₃ with theta and kappa A1 ₂ O ₃ porosity in cm ³ / g total radii 100 AE radii 100 to 500 AE radii 500 AE 0.44 0.0 0.13 0.33 0.52 0.03 0.09 0.40 0.35 0.07 0.15 0.13 specific surface area in m ² / g 12.0 11.0 32 distribution of active ingredients in carrier moldings peripheral depth intrusion about 0.6 mm peripheral depth intrusion about 0.8 mm evenness. throughout the carrier molding form of catalyst height in m diameter in mm tablets 4.70 ± 0.15 4.8 ± 0.2 tablets 4.75 ± 0.15 4.75 ± 0.45 tablets 3.4 ± 0.1 3.6 + 0.2

Example 4

The catalyst, characterized in Table 3, was charged to the experimental apparatus described in Example 1 for the selective hydrogenation of the technical unsaturated C2-fraction. The feed mixture contained 1,200 to 2,500 vol / ipfpm. carbon monoxide. The other process parameters were varied within the range shown in Example 3. The reactor was started according to Example 2. The experimental results are summarized in Table 5, the residual acetylene concentration in all cases being around 2 ppm.

Table 4

temperature on entry reactor Noc: 2 ° C coolant pressure Pa / cm ² load v / vh koncen- trace WHAT obj.- ppm balance C2H4 based on CaHi charged ±% 86 0,55.10® 5 800 6 300 + 0.08 76 0.50.10® 5 800 3 600 + 1,22 72 0,53.10® 5 800 2 500 + 0.61 83 0,85.10® 7 000 2 500 - 2,02 80 0,94.10® 4 400 2 400 - 2,55 80 0.95.10® 2 900 5 000 +1.05 95 2.5 5 800 2 600 - 5,46

Table 5

temperature reactor on entry Noc: 2 ° C coolant pressure Pa / cm ² load v / vh conc. WHAT obj.- ppm balance based on C ₂ H ₄ charged ± ±% 90 1,7 .10® 5 800 1900 + 1.24 90 1,6 .10® 5 800 1700 +1,5 90 1,5 .10® 5 800 1500 + 0.83 90 1,3 .10® 4 400 2100 + 1.19 87 1,2 .10® 2 900 2100 + 1.19 93 1,55.10® 7 000 1500 + 1.32 94 3,5 .10® 5 800 1400 - 0,24

Example 5

The catalyst, as described in Table 3, was charged to the experimental reactor of Example 1. The start-up was carried out analogously to Example 1, but with a carbon monoxide concentration of 1600 vol .- ppm. Another embodiment of the experiment corresponded to Examples 3 and 4. The results summarized in Table 6 are characteristic of this experiment. Acetylene was in all cases degraded to 3 ppm. The hydrogenation reaction is started at 57 ° C.

Table 6

temperature on entry reactor Noc: 2 ° C coolant pressure Pa / cm ² load v / vh conc. WHAT vol - ppm balance c ₂ h ₄ relative to the set C ₂ H ₄ ±% 72 0,35.10® 3 500 1600 + 0.62 76 0.80.10® 5 800 1800 + 0.94 85 0.95.10® 5 800 1200 + 0.78 86 1,5 .10® 5100 1200 - 0,52 85 2,5 .10® 5 100 1250 - 1,5 87 2,95.10® 5 800 1300 - 4,36

Example 6 - Comparative Example For comparison with the results contained in the Examples 1-5 of the invention were in the experimental apparatus used in Example 1 is charged with 7.2 1 and industrially produced in the selective hydrogenation of unsaturated _C2 -fractions be dried and the catalyst according to the principle used in Examples 1 to 5 and tested essentially under analogous conditions.

The comparison catalyst is characterized by the following data:

bulk density burst composition Pd

CuO Εθ2θ3 NiO

0.75 kg / l

356.10 ⁵ ± 290. ΙΟ ^{3 *}

Pa / cm ²

0.034 wt.

0.01 0.54> 0.015 carrier structure mixture consisting of alpha and theta alumina pore volume, total 0.60 cm ³ / g for r <100 A 0.25 cm ³ / g for 100 <r <500 A 0 25 cm ³ / g for r> 500 A specific surface shape of the catalyst

0.10 cm ^{1 2} / g 82 m ² / g tablets with a diameter of 6.0 + 0.8 mm with a height of 3.2 and 1.1 mm

Upon commissioning, a crude fraction having a carbon monoxide content of 3,000 ppm and a temperature of 25 ° C was charged to the reactor. The load was 5,800 v / vh. The inlet temperature was increased to 60 ° C in small increments over a period of 3 hours and maintained at that temperature for 1 hour until the reactor outlet temperature reached the maximum level. Then the temperature was further increased at a rate of 6 ° C / h. The hydrogenation reaction was started at 72 ° C. The desired degradation of acetylene to 3 ppm was achieved about 7.5 hours after the unsaturated C ₂ -fraction reactor was operated at an inlet temperature of 80 ° C and a coolant pressure of 0.8.10 ⁵ Pa / cm ² . The other parameters were varied within the range given in Example 1. The characteristic values summarized in Table 7 were obtained.

Starting a catalyst with a carbon monoxide concentration below 2000 ppm was not possible because it had a strong tendency to run. In addition, the process stability achieved compared to Examples 1 to 5 was considerably less. A relatively small increase in the reactor temperature (especially above the refrigerant pressure), or a decrease in the carbon monoxide concentration, led to too rapidly increasing losses of ethylene and ultimately even to the reactor run-out. This is particularly evident when comparing the results summarized in Table 8, given the stability of the process and the running of the catalysts.

Claims (15)

OBJECT OF THE INVENTION CLAIMS 1. A process for the selective hydrogenation of unsaturated C2-fractions containing, in addition to ethylene, ethane and methane, 0.3 to 2.0 wt. % of acetylene; hydrogen and 1000 to 2000 volume of the draw-ppm of CO at a pressure of 1.10 ⁵ to 40.10 ⁵ Pa, a space velocity of 2,500 to 10,000 v / vh and a reactor temperature of 35-150 ° C, characterized in that the tubular reactor filled with a catalyst having a concentration of carbon monoxide in the unsaturated C 2 -factor ranging between a process-dependent content of up to 5,000 v / ppm and a space velocity of at least 3,000 v / vh is put into operation using macroporous pores and optionally 0.0005 to 0.25 wt. copper, nickel, silver and / or iron or compounds of these elements, a supported palladium catalyst on a support having a specific surface area of 50 m ² / g or more and having a porosity of more than 500 AE with at least 33% porosity and a palladium concentration of 0,005 to 0; 1% by weight, wherein the flow rate of the unsaturated C2-fractions through the catalyst tube is between 0.3 to 3.0 m / s and the ratio of the clear tube width to the diameter or the height of the catalyst moldings is between 8 to 25 and the reactor inlet temperature is increased to near or below the catalyst start-up temperature over 1 to 6 hours, and then optionally, with intermediate interruptions of 30 to 60 minutes, at substantially constant inlet temperature, the desired coolant removal and temperature change of the reactor is carried out at a maximum rate of 20 ° C / h, the temperature difference between the outlets of the reactor and the reactor uptake and inlet to the reactor as well as between any point in the catalyst bed and the cooling medium is 0 to 60 ° C. 2. The process of claim 1 wherein the catalyst support consists essentially of alpha alumina. Process according to Claims 1 and 2, characterized in that a catalyst having a specific surface area of 3 to 15 m ² / g, preferably 3 to 10 m ² / g, of which at least 75% of the pore volume is attributed to pores with radii above 500 AE. 4. A process according to any one of claims 1 to 3, wherein the catalyst comprises from 0.01 to 0.05 wt. % palladium and as promoters 0.003 to 0.05 wt. % of nickel, copper and / or silver or their oxides, and 0.03 to 0.2 wt. ferric oxide or ferric oxide. Process according to Claims 1 to 4, characterized in that a catalyst with a palladium contained in the peripheral layer of the catalyst compact of at most 1.5 mm and preferably at most 1.0 mm is used. 6. A process according to claim 1, wherein the catalyst comprises from 0.005 to 0.05 wt. % of palladium on a support consisting essentially of alpha alumina with a specific surface area of 15 to 50 m ² / g, with at least 33% of the pore volume attributable to pores with radii above 500 AE. 7. The method of claims 1 to 6, wherein the ratio of the internal diameter of the catalyst tube to the ratio or height of the catalyst moldings is between 10 and 15. 8. Process according to claims 1-7, characterized in that the flow speed unsaturated _C2 the catalyst tube fraction is 0.5 to 2.0 m / s. 9. The process of claims 1-8 wherein the temperature at any point in the catalyst bed is at most 45 [deg.] C above the temperature of the cooling medium. 10. The process of claims 1 to 9 wherein the difference between the reactor inlet temperature and the reactor inlet temperature is between +50 to -f-5 ° C. 11. The process of claims 1 to 10, wherein the reactor inlet temperature is maintained at up to 20 [deg.] C. below the catalyst start-up temperature, wherein the start-up temperature is the temperature at which the catalyst just begins to hydrogenate acetylene under given process conditions. 12. A process according to any one of claims 1 to 11, wherein the concentration of carbon monoxide in the crude fraction is between the process conditioning concentration and a maximum of 3500 vol. is at least 3,000 v / vh, preferably 3,500 to 4,000 v / vh, the inlet temperature of the reactor is increased in small steps up to or below the catalyst feed temperature in 1 to 6 hours, preferably 2 to 3 hours temperature. 13. The process of claim 1, wherein the concentration of carbon monoxide in the crude fraction is between the process conditioning content and 3000 vol .- ppm when the catalyst is reactivated. Process according to one of Claims 1 to 13, characterized in that the inlet temperature of the reactor varies in small increments, but generally at most 20 ° C / h, preferably 10 ° C / h. Method according to one of Claims 1 to 14, characterized in that the temperature of the cooling medium varies in small increments, but generally by a maximum of 10 ° C / h and preferably by a maximum of 5 ° C / h.