CA1053265A

CA1053265A - Process for the production of hydrocarbon mixtures

Info

Publication number: CA1053265A
Application number: CA250,791A
Authority: CA
Inventors: Walter Rottig
Original assignee: Ruhrchemie AG
Current assignee: Ruhrchemie AG
Priority date: 1975-04-29
Filing date: 1976-04-22
Publication date: 1979-04-24
Also published as: JPS5443481B2; AU1322476A; DE2518982B2; AU498400B2; NL162622C; FR2309497B1; ZA762412B; JPS51131808A; GB1490494A; NL162622B; BE840969A; FR2309497A1; NL7509662A; DE2518982A1; DD123881A5

Abstract

ABSTRACT OF THE DISCLOSURE

There is disclosed a method for the production of hydrocarbon mixtures containing at least 50% by weight of C2 to C4 hydrocarbons and at least 50% by weight of olefins by catalytic hydrogenation of carbon monoxide in fixed-bed re-actors under certain particular conditions which are necessary for the production of the aforementioned hydrocarbons.

Description

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The catalytic hydrogenation of the oxides of carbon is broadly known. These prior art hydrogenation reactions pro-duce mixtures of paraffins and olefins containing 1 to 40 carbon atoms and, in many cases, also produce oxygen containing com-pounds such as alcohols, aldehydes, ketones, esters, or fatty-acids. Minor proportions of aromatic hydrocarbons are also pro-duced under selected synthesis conditions ~see Ullmann, Encyclo-padie der technischen Chemie. 1957, Vol. 9, pp. 701 et seq.).
It ha~ also been well-known that the elements of Group VIII of the Periodic Table are exceedingly active for the hydrogenation of the oxides of carbon. In particular, iron, cobalt, and nickel are useful in this reaction. Due to their high hydrogenation activity, nickel and cobalt catalysts yield predominantly mixtures o saturated, straight chain hydrocarbons.
On the other hand9 such catalysts based on iron permit the pro-duction of hydrocarbon mixtures containing high proportions of unsaturated aliphatic compoundsc In addition, the iron base catalysts produce oxygen containing compound~ especially ali-phatic primary alcohols (UlLmann, op.cit., pp. 697-698).

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It is also easy to produce mixtures containing at least 60% by weight o~ hydrocarbons having more than about 20 carbon atoms. Such hydrocarbons would correspond to an atmo-spheric boiling range in excess of about 320C (see Ullmànn, op.cit., pp. 722~.
However, prior to ~he present invention, it was not possible to direct the hydrogenation reaction of the oxides of carbon to reaction products which contain more than 50% ~(~y weight) of low molecular weight (preferably olefinic) hydro-carbons having 2 to 4 carbon atoms The percentage set forth is based upon the total hydrocarbons having at least 2 carbon atoms.
It is therefore an object of this invention to pro-vide a process for the preferential production of low molecular weight olefinic hydrocarbons by the catalytic reduction of oxides of carbon in the presence of hydrogen.
It has now been found that it is possible to produce hydrocarbon mixtures containing at least 50% by weight of C2 to C4 hydrocarbons and at least 50% by weight of olefins, all based on the total content of hydrocarbons having at least two carbon atoms, by the ca~alytic hydrogenation of the oxides of ~ -carbon in fixed-bed reactors under elevated tempera~ure and pres- , sure. These results can be achieved in a process in whiclh the catalyst is arranged in at least 2 and not more than 10 layers in reactorshaving a length of about 0.5 to 4.5 ~eters. The synthesis pressure is about 5 to 30 bars, and the synthesis temperature is about 250 to 370C., measured on the catalyst.
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The fresh gas load (or space veloci~y) is adj~ ted to about 1,000 to about 10,000 standard cubic meters per cubic mcter of catalyst per hour. The gas recycle must be about 5 to 25 times the amount of fresh gas and the ~o~al gas load must be from about 7,500 to about 50,000 standard cubic meters per cubic mcter of catalys~ per hour. The catalys~ is arranged so that the linear gas velocity is between about 1 and 10 meters per second, based "
on standard conditions. The residence time is between O.OS and 1 second, based on standard conditions, and the catalyst used has an internal surface area of,about 5 to 150 square meters per gram.
The catalysts are arranged in a fixed-bed in 2 to 10 ,;
layers. They may have the same or differentgrain sizes, but it is desirable that they be about 0.5 to about 10 millimeters and most preferred about 1 to 6 millimeters. The specific ~in size - is dependent upon various factors such as the activity of the catalyst, the diameter of the reactor, the resistance to flow and the operating conditions.
The shape of the catalyst is of minor importance.
Cylindrical pellets, spherical pelletsS tablets~ polyhedra, and hollow bodies as well as lumpy, crushed and/or comminuted catalysts may all be used with good results.
The reactor may contain tubes having an inside diameter of, for example, a minimum of about 30 millimeters. These tubes may also be in bundles, or vessels such as shaft furnaces or similar apparatus may be used as fixed-bed reactorsO The reactors may be provided with devices for cooling or removing the heat of reaction.
In that even~J cooling media such as water, liquid salt melts, or organic cooling liquids (based, for exampIe, on biphenyl, o-dichlorobenzene or isomeric benzyl benzenes) may be used in the case of the tubes Or tube bundles.

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~053~;5 It is particularly advantageous to remove the heat o~ reaction almost completely by recycling the gas mixture. If necessary or desired, a cooling medium can also be used. When operating in this manner, the temperature in the catalyst bed is not uniform throughout, but increases in a con- ¦
trolled manner in the direction of the ~low of the gas stream.
Thus, the gas mixture enters the reactor at a temperature of, for example, 280C. This temperature may increase as the gas r flows through the reactor b~ about 50C. to a temperature of 330C. This is the result of absorption of part of the heat of reaction by the gas mixture. ~he temperature of the catalyst itself increases by approximately the same a~ount and the over-all gas mixture leaves the reactor at this higher temperature.
This increase in temperature (the vertical temperature gradient) should be at least 25C but should not exceed 100C. ~he pre-i .
ferred vertical temperature gradient is 30 to 50C. r:In the preferred form of the reaction, the heat of F
reaction is removed by appropriately adjusting the gas recycle, without the use of a cooling medium. For teck~cal and economic ~-reasons, it is advisable to operate in a plurality of stages and to cool the reactio~ mixture partially or completely between the stages. If desired, the reaction products may be separated ~rom the reaction li~uor after each stage. The number o~ stages is approximatel~ 2 to 10 and preferably 2 to 5.
In a variation of the present invention, cold fresh ~-gas may be fed down stream of each of the aforementioned stages r to partially cool the reaction mixture an~ to increase the pro-portion of carbon monoxide and hydrogen in the gas mixture. It is also possible, in this form of the reaction, to increase the reaction temperature from stage to stage to permit better con- r .
, '' ~ _ 16~5;3~j5 trol of the conversion of the carbon monoxide and hydrogen.
This con~ersion decreases as the content of the starting materials in the reaction mixture goes down. This occurs, naturally, as a result of these materials being used up in the reaction itself. It has been found useful to increase the reaction temperature by about 5 to 20C from stage to stage.
The length of the reactors used to carry out the reaction must not exceed 4.5 meters. It is preferred to use reactors having a length of less than 3 meters. Re-actors as short as 0.5 to 1.0 meters can be used without disadvantages in the present reaction.
The pressure used is dependent upon the na~ure of the catalyst and ranges from 5 to 30 bars. More preferred are pressures between 5 and 1~ bars. It has been noted that lower pressures frequently result in a decrease in the con-version of the starting materials.
The reaction temperatures are between 250 and 370C., preferrably between 270 and 340C.
It is of substantial importance for the proper operation of the present process that a high fresh gas load be maintained.
In prior art fixed-bed catalyst reactions, a load of 500 to 700 standard cubic meters per cubic meter of catalyst per hour was generally considered to be the upper limit in the Fischer-Tropsch synthesis. Fresh gas rates greatly in excess of this "limit" are needed in the present invention. More specifically, it has been . ~.
ound desirable to supply 1,000 to 10,000 standard cubic meters , of fresh gas per cubic meter of catalyst per hour. Preferably, . ~ , . . : .
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l,S00 to 5,0~0 standard cubic me~ers of catalyst per hour are used, In a~dition to tl-e fresh gas load, ~he invention re-quires that the(g-as be recycled. It is of particular iMportance that the ratio of ~he recycled gas to the fresh gas be maintained within certain limits. The recycled gas may be conducted through a plurality o~ stages before recycling but must be adjusted to about 5 to ~S times the amount of fresh gas introduced. It is most preferred that the ratio of recycled gas to fresh gas be from 7.5 to 15. In addition to ~he foregoing, the total gas load to the catalyst (the sum of the fresh gas and the recycle gas) must be between about 7,500 and about 50,000 standard cubic meters per cubic meter of catalyst per hour. The preferred total gas load being about 10,000 to 25,000 standard cubic meters of gas per cubic meter of catalyst per hour.
The linear flow velocity of the gas mixture and the re-sidence time in the catalyst must also be controlled. Both of these variables are based on standard conditions and the linear gas flow velocity must be from 1 to 10 meters per second. The preferable range is 1.5 to 5.0 meters per second. The residence time is 0.05 to l second and the preferable range is from 0.05 to 0,5 seconds. I-t has been noted that the production of hydro-carbons of higher molecular weights is promoted by an increase in residence time.
The catalysts useful in the present invention are those normally used in the Fischer-Tropsch synthesis. These are generally known and are operable in the present process. However, it has been found that catalysts containing more than 50V/o by weight of iron are to be preferred. More preferred, are catalysts containing more - -, r~?_ _ _ 1~ 5 ~ 6 5 than 60% by weight of iron. In addition, promoters such as copperand/or silv~r and alkali are desirable. Moreover, o~her additives such as alkaline earth metal compounds, zinc oxide, manganese oxide, ceri~m oxide, vanadium oxide, chromium oxide and the like may also be used. On the other hand, the use of support materi~ls such as alumina, kieselguhr, or impregnating agents such as potassium or sodium water glass is less adv~ntageous.
Boron, phosphorous, tungsten or molybdenum may be used as components of the catalyst in the form of their oxides or salts. Such salts as sodium borate or potassium tungstenate are recommended. It has been found particularly desirable to use catalysts based on iron, copper and/or silver and alkali (K20).
More specifically, such a catalyst in a ratio of Fe:Cu/Ag:K20 of, for example, 100:3 to 25:10 has been found especially suitable.
The catalysts themselves may be produced in any known manner; for example, precipitation, sintering, fusion or de-composition of salt mixtures. Shaping, reductlon, etc., of the catalysts may also be carried out in the known manner. In some cases, sintered catalysts have been found to be somewhat more i~ ~ advantageous.
The internal surface area of the catalysts is measured by the BET method and is important for the proper operation of . , the process of the present invention. Catalys~ having an internal ; surface æea in excess of 150 square meters per gram are likely to produce higher molecular weight hydrocarbons. Therefore, the catalysts used in the present process must have an internal sur-face area of about 5 to about 150 square meters per gram of cata-.
, lyst. It Ls pre~erred that this range be about 10 to about 100 - , . . . .
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l~S~;5 square meters per gram of catalyst. For those catalysts which contain predominant amounts of me-tals, especially ironj the internal surface area is bases on the reduced state. ~-~r ~
It has been found tha-t the pore size of th~catalysts r has a noticable influence on the proportion of low molecular ¦ -weight hydrocarbons in the reaction product. The proportion of macropores having a diameter of more than about 5 x 10 6 to 1 x 10 5 centimeters bases on the total pore volume should be as low as possible since the macropores promote the production of the higher molecular weight hydrocarbons. Pores having a diameter of less than 5 x 10 6 centimeters encourage the pro- ;
duction of the lower molecular weight hydrocarbons. It i-s, therefore, advantageous if the catalysts used in the present process contain less than 50 % macropores b s on the total pore volumeO It is to be preferred that the catalysts contain less than 25 % of the macropores - It occasionally happens during the course of operation o~ the present process, that some of the catalyst pores become clogged by~deposition of small amounts of high molecular weight .~
reaction products. When -this occurs, portions of -the catalysts ^
are no longer accessible to the synthesis gas. ~his condition ~evidences itself by a decrease in the conversion of carbon monoxide and hydrogen. The high molecular weight deposits may be removed by what is known as an extractive operation, i.e. ex-trac-tion of the catalyst with hydrocarbon mixtures produced by the synthesis itself or with other hydrocarbon mixtures. r .
The catalysts may be produced in a manner kno~m per se, I
e.g. by precipitation, sintering, fusion or decompasition of salt mixtures. Molding and reduction of the catalysts may also be effected in kno~m manner. In some cases, sintered catalysts have been found to be advantageous r The ratio of carbon monoxide to hydrogen in bhe fresh gas also has an influence on the production of low molecular 8 ~ ~

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weigh~ hydrocarbons. While it is possible, in the present process, to use gases which are rich in carbon rlonoxide, it is preferred that the fresh gas be rich in hydrogen. The rati~ of carbon monoxide to hydrogen in the fresh gas is preferably in excess of 1:1.2. It is most desired that the ratio be between 1:1.5 and 1:2 Higher proportions of hydrogen may be somewhat dis-advantageous under cer~ain circumstances.
The presence of inert gases such as methane, carbon dioxide, or nitrogen, generally does not interfe~ with the re-action. Since the proportion of these inert gases in the reaction mixture increases as the conversion of carbon-monoxide and hydrogen increases due to the gas contraction which occurs, it is desirable to keep the inert gas concentration in the fresh gas low.
If the process is practiced in accordance with the pre-sent invention, it will yield more than 50% by weight and, in many cases, more than 60% by weight of hydrocarbons having 2 to 4 carbon atoms, based on the total amount of hydrocarbons excluding methane. The proportion of olefins is in excess of 50%, also based on the total hydrocarbons with the exclusion of methane.

~ A catalyst in the form of spherical pellets Sabout : 2 to 2 . 5 millimeters in diameter) is prepared by sintering a homogeneous mixture of iron oxide (Alan Wood ore), copper oxide, zinc oxide, and potassium carbonate in the ratio of Fe:Cu:ZnO:K20 of 100:25:10;8 for 2 hour~ at 1050C, The catalyst was reduced for several hours at 400C. After termination of the treatment, the catalyst had a reduction value (percentage of free iron based _ g _ ~ , :

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' 105;~;~t,i5 on total iron) of 97 %. ~he catalyst was placed in an electrically heated test reactor 1 meter in length and 50 millimeters in inside diameter. It formed a layer 50 centimeters in depth.
Fresh gas having the following composition:
C2 3.6 ~o by volume ~
CnHm O % by volume n~,2 m-2n _ 2 0 % by volume CO ~0.7 % by volume H2 53.0 % by volume ~
CH4 0.2 % by volume _ N2 12.5 % by volume C0/H2 ratio 1 : 1~72 was introduced into the reactor at a rate of 2,000 standard liters ¦
per liter o~ catalyst per hour (2,000 v/v/hr.), Recycled gas ~`
was introduced at the rate of 15,000 standard liters per liter F.:`
of catalyst per hour (15,000 v/v/hr.). The pressure was maintained at 10 atmospheres. A conversion of the carbon monoxide and hydrogen of 36 ~o was obtained at a temperature of 260C measured outside of the catalyst, and at a temperature t of 292C measured inside the catalyst.
~ he reaction mixture had the ~ollowing composition based on hydrocarbons havin~ 2 or more carbon atoms:
C /C about 69.5 ~o by weight; ~;
2 4 of these, 5? % by weight r-were olefins;
C to about C about 30.4 % by weight;
11 of these, 60 % by weight were olefins. L
~he proportion of methane was between 7 and 8 ~ by ~reight.

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The sintered catalyst was produced in a~cordance with Example 1 except that only 5 parts 1~y ~eigh~ of copper and 4 parts being weight of K20 were used. Othe~ise, the conditions and composition were unchanged. The fresh gas and recycle gas rates ~ere the same as in Example 1, but the reaction temperature was 270~C outside of the catalyst and 313C inside the catalyst. As a resul~, the carb~n monoxide and hydrogen conversion rate increased to about 40%. The composition of the synthesis products ~as as follows:
C2/C4 about 98% by weight ~ C4 about 2% by weight The olefin content of the C2/C4 fraction was 56~/o by weight. The proportion of C2 hydrocarbons in the C2 /C4 fraction was about 90%, the balance being about equal amounts of C3 and C4 hydrocarbons. Methane in an amount of 10 to 12%
; by weight was also produced.
The synthesis gas used had the sanle composition as that set forth in Example 1. The carbon monoxide/hydrogen ratio was 1:1.3 rather than 1:1.72 in the synthesis gas, and the proportion of olefins increased by about 4 to 5~0 by weight.
~hile only a limited number of embodiments of the present process have been specifically described, the invention is, nonetheless~ to be broadly construed and n~ to be limited except by the character of the claims appended hereto.

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i~5;~'~6~i SUPPLEl`1ENTARY DISCLOSURE
_ The principal disclosure describes a full procedure for producing hydrocarbon mixtures containing at least 50% by weight of C2 to C4 hydrocarbons and at least 50% by weight of olefins, all based on the total content of hydrocarbons having at least two carbon atoms, by the catalytlc hydrogenation of the oxides of carbon in fixed-bed reactors under elevated temperature and pres- ~ -sure. The following examples represent further specific procedures carried out within the scope of the ir~vention claimed in the original disclosure.
Example 3 The reaction mixture obtained in Example I of the principal disclosure was further processed in the ` follo~ing manner: ~
The reaction mixture was conducted under pressure -over activated carbon to remove the hydrocarbons which are '~ gaseous standard conditions. The residual gas had the foll-` owing composition:
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~!~ C O . 22.8 ~ H2 48.8 "

" CnH2n+2 1.4 N2 16.7 which was then led to a second reaction step identical with the first. The operating conditions were the same except j for the temperature which was 270C outside the catalyst and 291C inside the catalyst. A C0/H2 conversion rate of about 34% was obtained.

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C atoms was as follows C /C 7 The proportion of newly formed methane was about wt. , based on all the hydrocarbons formed while the catalys~ used corresponded to that of step l.
Exa~ple 4 The product obtained in Example 2 of the principal . dlsclosure was further processed in the following manner:
After separation of the gaseous hydrocarbons under standard conditions, by means of activated carbon under pressure, the residual gas had the following composition:
C2 10.2 vol. %

i ; CnHm - 0 4 " :
2 o C0 22.1: "
H2 48.0 ~20~. CnH2n+2 l.g -N2 17.4 ll : This gas was then reacted in~a second reaction ste comparable:to the first, and-under identical operating conditions except for~the reaction temperatures of 280C outside the catalyst and 310'C inside the:catalyst. A 37% conversion of ;carbon~monoxide and hydrogen was obtained :;Th:e~co~mposit~ion Or the~reuulting products with more than:~-two;~oarbon~atoms was~as r~llows;
C2/~C4 ~ : about 97 wt.Yo 4 ~ about ~ ll he percenta~e of~the G2 rraction in the C2/C
h~rocarbons was~:~about û5~:wt.%~while 8 wt.Yo conuisted Or lOS~ iS

C3-, and 7 ~.% o~ C4-hydrocarbons. The olfein content o~
the C2-hydrocarbons was 53 wt.,', ~hile the o~in content of the C3-~raction ~ras about 70 wt.% and a~out 65 wt.Y for the C4-fraction.
~ he methane formation ~as about 13 ~/o and the catalyst used corresponded to that o~ step 1.

~X~ ~LE 5 Iron oxide (Alan Wood ore) was finely mixed with some copper oxide, zinc oxide and potassium carbonate and formed on a rotary plate to produce an spherical catalys~ (2 to 3 mm diameter). ~he composition had a ratio of: 100 Fe: 2 Cu:
4 ZnO: 2 K20. ~he catalyst ~ras sintered for three hours at 1100C. and, after cooling reduced ~rith hydrogen for several hours at 430C.
Ten adjacent, serially connected, electrically heated reactors (clear width 60 mm, height 1.25 m) were filled with the above-prepared catalyst. The first stage contained 3.55 liters of catalyst. The amount of catalyst was decreased from stage to stage by about 5 percent by volume so that the tenth and last stage contained 1.8 ~lters of the catalyst.
After removal of the reaction gas from the first sta~e and subsequent cooling, the resulting C2/G4 hydrocarbons was separated from the reaction gas by passing said gas through an activated carbon layer under pressure. The hydrocarbons were removed from said activated carbon layer by contact with water vapor at a temperature of approximately 100C. and about standard pressure by collecting said gaseous mixture in a gasometer. The remaining gas was then transferred to the second stage.
After passing into the second reaction and subsequent cooling, the recovery of the newly formed C2/C4 hydrocarbons C

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took place as previously described. This process was carrled out in the same manner for the third through the ninth reaction stage. The tenth and last reaction stage did not have the remain-ing gases pass through an activated carbon layer. Instead, the recovery of the C2/C4 hydrocarbons took place by means of cooling at a temperature of minus 60C. wherein condensation of the desired hydrocarbons was effected.
The reaction temperatures, measured inside the catalyst, were between about 310C. (first stage) and about 345C. (last stage). Thus, the reaction temperature was increased from stage to stage by about 5C.
Each stage was provided with recycled gas. The quan-tity of the recycled gas amounted to about six times the feed gas entering the single reactor. 2,500 standard liters per liter of catalyst per hour (2,500 v/vhr) o~ a synthesis gas with the following composition were passed from the first stage at a pressure of 11 bars:

% by volu~e The carbon monoxide and hydrogen conversion in each stage amounted to about 20 percent based on the portion of these gases present in the inlet gas before each stage and after separation of the C2/C4 hydrocarbons.
Based on the procedure of the foregoing example, a total carbon monoxide and hydrogen conversion rate of 89.1%

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was attained based on a ten s~age reaction. The composition of the thus formed synthesis products per Nm3 ~f fed synthesis gas was as follows:
CH4 15.8 grams C2 hydrocarbons 145 . 2 grams - 55% by weight were olefins C3 hydrocarbons 6.4.grams - 65% by weight were olefins C4 hydrocarbons 8.0 grams - 60% by weight were olefins :
C5 and higher hydrocarbons trace amounts ~ . . .

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Claims

The embodiments of the invention in which an exclusive property or privilege is claimed are defined as follows:

1. In a process for the production of hydrocarbon mixtures by catalytic conversion of gas comprising carbon mon-oxide and hydrogen in fixed bed reactors at elevated temperatures and under superatmospheric pressure, the improvement which com-prises, (a) arranging the catalyst in from 2-10 layers in said reactors, (b) said reactors having a length of 0.5 to 4.5 meters, (c) said pressure being 5 to 30 bars, (d) said temperature being 250° to 370°C measured in said catalyst (e) introducing fresh gas at a space velocity of about 1,000 to about 10,000 standard m3 per m3 of said catalyst per hour, (f) recycling gas after contact with said catalyst, said recycle gas being introduced at a rate of about 5 to 25 times said space velocity, (g) said fresh gas and said recycle gas constituting a total gas load, said total gas load being about 7,500 to about 50,000 standard m3 per m3 of said catalyst per hour, (h) the linear velocity of said gas being from about 1 to 10 meters per second, based on standard con-ditions, (i) the residence time of said total gas load in said catalyst being 0.05 to 1 second, based on standard conditions, and (j) said catalyst having an internal surface area of about 5 to 150 m2 per gram of said catalyst, whereby said mixtures contain at least 50% by weight of C2 to C4 hydrocarbons and at least 50% by weight of olefins, based in each case on the total hydrocarbons present having at least 2 carbon atoms.

2. A process according to claim 1 wherein the length of said reactor is less than 3 meters.

3. A process according to claim 1 wherein said pres-sure is 5 to 15 bars.

4. A process according to claim 1 wherein said tempera-ture is 270 to 340°C.

5. A process according to claim 1 wherein said space velocity is 1,500 to 5,000 standard m3 of catalyst per hour.

6. A process according to claim 1 wherein the rate of introduction of said recycle gas is 7.5 to 15 times said space velocity.

7. A process according to claim 1 wherein said total gas load is 10,000 to 25,000 standard m3 per m3 of catalyst per hour.

8. A process according to claim 1 wherein said linear gas flow velocity is from 1.5 to 5 meters per second.

9. A process according to claim 1 wherein said resi-dence time is from 0.05 to 0.5 seconds.

10. A process according to claim 1 wherein said catalyst has an internal surface area of 10 to 100 m2 per gram of said catalyst.

11. A process according to claim 1 wherein said catalyst has an iron content of more than 50% by weight.

12. A process according to claim 11 wherein said catalyst contains more than 60% iron by weight.

13. A process according to claim 1 wherein said catalyst contains less than 50% macropores having a diameter of more than about 5 x 10-6 to 1 x 10-5 cm.

14. A process according to claim 13 wherein said catalyst contains less than 25% of said macropores.

15. A process according to claim 1 wherein said catalyst is in 2 to 5 layers.

16. A process according to claim 1 wherein addi-tional CO and H2 is introduced to said gas load downstream of the point at which said recycle gas joins said fresh gas.

17. A process according to claim 1 wherein said gas is rich in H2.

18. A process according to claim 1 wherein the ratio of CO/H2 in said gas is at least 1:1.2.

19. A process according to claim 18 wherein said ratio is from 1:1.5 to 1:1.2.

20. A process according to claim 1 wherein the catalyst has pores less than 5 x 10-6 cm in diameter.