Title of Invention

PROCESS FOR THE RECOVERY OF SULPHUR FROM HYDROGEN SULPHIDE CONTAINING GAS

Abstract The invention relates to a process for the recovery of sulfur from a hydrogen sulfide containing gas, comprising (a) Oxidizing part of the hydrogen sulfide in a gaseous stream with oxygen or an oxygen containing gas in an oxidation stage to sulfur dioxide, and thereafter reacting the major part of the remaining hydrogen sulfide and the major part of sulfur dioxide to elemental sulfur and water, (b) reacting the product gas of oxidation stage a) in at least three catalytic reaction systems and c. selectively oxidizing H2S in the gas leaving stage b (in particular leaving) the (last) sub-dewpoint Claus reactor) to elemental sulfur, preferably employing for this purpose a catalytic stage including a selective oxidation catalyst, which is substantially insensitive to the presence of water vapor in the gas stream and is ineffective in promoting the establishment of the Claus equilibrium: 2 H2S + SO2 2 H2O + 3/n Sn.
Full Text

Title: Sub-dewpoint Claus and hydrogen sulfide oxidation
In a number of processes, such as the refining of crude oil, the purification of natural gas and the production, of synthesis gas from, for example, fossil fuels, sulfur containing gas, in particular HsS containing gas, is released. On account of its high toxicity and its smeil, the emission of H2S is not permissible.
The best-known and most suitable process for regiovrag sulfur from gas by recovering sulfur from hydrogen sulfide is the so-called CT&us process. In this process hydrogen sulfide is converted by oxidation to a considerable extent into elemental sulfur; the sulfur thus obtained is separated from the gas by condensation. The residual gas stream (the so-called Claus residual gas) still contains some HsS and SO2.
The method of recovering sulfur from sulfur containing gases by the so-called Claus process is based on the following overall reactions:

A conventional Claus plant - suitable for processing gases having axi H2S content of between 50 and 100% - comprises a burner with a combustion chamber, the so-called thermal stage, followed by a number of reactors -generally two or three - filled with a catalyst. These last stages constitute the so-called catalytic stages.

In. the combustion chamber, the incoming gas stream, which is rich in H2S, is combusted with an amount of air at a temperature of Approximately 1200°C. This amount of air is adjusted so that one third of the H&S is fully combusted to form SO2 in accordance with the following reaction;
2H2S + 3O2 => 2H2O + 2SO2 (1)
After this partial oxidation, of H2S the non-oxidized part of the H2S (i.e. basically two-thirds of the amount offered) and the SO2 formed react further as to a considerable extent, in accordance with the Claus reaction:
4 H2S + 2 SO2 4H2O + 3 S2 (2)
Thus, in the thermal stage, approximately 60-75% of the H2S is converted into elemental sulfur.
The gases coming from the combustion chamber axe coaled to about 160°C in a sulfur condenser, in which the sulfur formed is condensed. Subsequently, the condensed sulfur flows into a sulfur pit through a siphon*
The non-condensed gases, of which the molar ratio of IfeSrSOa is unchanged and still 2:1, are subsequently heated to about 250°C> and passed through a first catalytic reactor in which the following equilibrium is established;
4 HsS + 2 SO2 «» 4 BtO + 6/ft Sn (2)
The gases coming from this catalytic reactor are subsequently cooled again in a sulfur condenser, from which the liquid sulfur formed is recovered and the remaining gases, after being reheated, are passed to a second catalytic reactor.

If the gaseous feedstock contains H2S concentrations of about 15 to 50%, the above described "straight-through1' process is not used, but instead a variant thereof, the so-called "split-flow" process. In the latter process one-third of the total amount of feedstock is passed to the thermal stage and combusted completely to SOs therein. Two-thirds of the feedstock is passed directly to the first catalytic reactor, bypassing the thermal stage. When the feedstock contains H2S concentrations of less than 15% the Claus process can no longer be used, The process then used is, for example, the so-called Recycle Selector process, in which the feedstock is passed with an adjusted amount of air into an oxidation reactor, the so-called oxidation stage. The reactor contains a catalyst, which promotes the oxidation of H2S to SO& and the amount of oxidation aix is adjusted so that an HsS:SO2 ratio of 2:1 is established, after which the Claus reaction proceeds. The gas forfeit the oxidation reactor is cooled in a sulfur condenser, in which the sulfur formed is condensed and discharged-
To dissipate the reaction heat generated in the oxidation reactor, a portion of the gas stream comiag from the sulfur condenser is recticulated to the oxidation reactor*
It is clear that in the Recycle Selected process, the oxidation, stage, which is catalytic and does not lead to high temperatures, is equivalent to the thermal stage in the Claua process. In the following, both the thermal Claus stage and the oxidation stage of the Recycle Seleetox process are deferred to as oxidation stages.
The sulfur recovery percentage in a conventional Claus process configuration is 92-97%, depending on the number of catalytic stages.
By known processes, the H2S present in the residual gets from the Claus converter is converted, by combustion, or some other form of oradation, into SOa, where after this SOa is emitted to the atmosphere. This has been permissible for low concentrations or small amounts of emitted St>2 for a long

time. Although 3O& has less noticeable impact than HgS, it is so harmful that its emission is limited by ever-stricter environmental legislation..
In the Glaus process as described above, in view of the equilibrium reaction which occurs, the BfeSiSCte ratio plays an important rote. In order to obtain an optimum, conversion to suliur, this ratio should be 2:1. In general, this ratio is controlled by means of a so-called H2S/SO2 residual gas analyzer. This analyzer measures the H2S and SO2 concentrations in the residual gas. A controller then maintains the ratio of 2:1 constant on the basis of the equation:
[HsS] - 2 [SOd = 0,
by varying the amount of combustion air, depending oft the fluctuations in the feed gas composition and the resulting deviation in the above equation. Such a control of the process, however, is highly Sensitive to these fluctuations, resulting in often less than optimal conversion
A second negative factor is that the theoretical sulfur itecovery efficiency (calculated on the basis of the amount of HgS supplied) is limited to approx. 97%, due to thermodynamic constraints imposed on the C/laus equilibrium reaction. As a result, the residual, gas from the last catalytic stage still contains substantial amounts of H2S and SO2 (in a molar ratio of 2 : 1).
The H3S present in the residual gas can be separated by absorption in a liquid.
The presence of SO2 in the residual gas, however, is a clisturbmg_ factor in the H2S separation process. It must therefore be removed first, complicating the residual gas treatment.
The SO2 in the residual gas reacts with conventional liquid HsS absorbents, forming undesirable products. To prevent undesirable reactions, the SO2 is generally catalytically reduced with hydrogen to form H2S over an AI2O3 supported cobalt-molybdenum catalyst, in accordance with the so-called

SCOT process.
The total amount of H2S is subsequently separated by liquid absorption in the usual manner.
In the SCOT process the sulfur components, other that H28, such as SOa (sulfux dioxide) and sulfur vapor (Se and Ss) are fully hydrogenated to H2S according to the following reactions:

Other components, such as CO, COS and CSa, are hydrolyzed according to:

Above conversions to H2S are performed with a cobalt-molybdenum catalyst on alumina at a temperature of about 280-330°C, For the SCOT process it is required that sulfur vapor is hydrogenated to IfeS, axfcd also that SO2 is completely converted to HsS down to ppm level, to prevent plugging/corrosion in the downstream water quench column. This type of hydrogenation can be defined as high temperature hydrogenation
In accordance with another method, for example, the BSR Selector process, after reduction of the SO2 in the residual gas to H2S and aifcer condensation of the water vapor, the gas is passed into an oxidation reactor, as in the Recycle Selector process. The oxidation, air is adjusted so that an HsSiSOa ratio of 2:1 is obtained, where after the Claue reaction proceeds. Both

in the SCOT process and in the BSR Selected process, the removal of SOa from the residual gas is a relatively expensive operation.
The above described after-treatment of the gases, carried out by means of a so-called tail gas treated, which involves an investment of another 50-100% of the cost of the preceding Claus converter, can result in an increase of the sulfur recovery efficiency of up to 98-99.8%.
Another group of processes to increase the sulfur recovery efficiency of a Claus process axe the so-called sub-dewpoint processes, suck as CBA (Cold Bed Absorption), MCRO (Maximum Claus Recovery Concept), S^cilfreen and Maxisulf. In these processes, one or more catalytic reactors are operated below the sulfur vapor dewpoint By doing this, the exothermic Claus Equilibrium reaction
2 H2S + SO* ^ 3/n Sn + 2 H2O (2a)
shifts to the right-hand side because of the lower catalyst temperature, and shifts even further to the right-hand side, because the major part of the produced elemental sulfur vapor Sn is removed from the gas phase by condensation into the pores of the Claus catalyst. The majority of the sulfur vapor produced by the Claus reaction is condensed in this way itf. the catalytic bed, and the remaining sulfur vapor corresponds with, and is limited to, the sulfur vapor pressure at the obtained catalyst temperature.
Consequently, the so-called sub-dewpoint reacfcor(s) art; slowly loaded with liquid sulfur in the catalyst pores- When the pores a.te almost completely filled with liquid sulfur, the Claus reaction will not proceed anymore, and the catalytic sub-dewpoint bed has to be freed frorft liquid sulfur by means of evaporation. This is indicated as the regeneration of the sub-dewpoint reactor.
The above-mentioned processes differ from each other in the this regeneration is performed. For CBA, the hot process gas froth the

preceding conventional catalytic Clans reactor is used for this ptirpose. For
MCRC, the hatjpxoe^^ data reactor is
cooted in. a sulfur condenser, after which it is reheated in a re-heater to 300-350°C and directed to the sub-dewpoiittjceactQi: to be regenerated.
The last reactor is always a sub-dewpoint reactor opetf&ted at a low temperature of typically 125-150°C. In the Stdireen process, the sub-dewpoint reactor to be regenerated is blocked-in into a closed recycle loop, which contains a recycle blowerr a reheater and a sulfur condenser,
In the sub-dewpoint processes, sxilfux is formed by the Claus reaction of H2S and SO2 according to reaction (2a). For optimum sulfur recovery efficiency, the ratio H;&S:SO2 should be 2:1 as for a "norrRial" Claus process. This means, that if excess HsS is available compared to SO2 (ratio H2S:SO2 > 2:1), the excess HsS cannot be converted to elemental stdfur^because there is not enough SO2 avaflable-
Consequently, the sulfur recovery efficiency will decease. Excess H2S will occur if not enough combustion air is supplied to the main (H2S) burner at the front end of the sulfur recovery unit. The same decrease in recovery efficiency will occur when SO2 is present in excess of H2S (ratio HsS:SO2 which is attainable for this type of processes. Consequently, the -correct- amount of combustion air is a crucial process parameter. The amount of combustion air, however, is difficult to control. This is a major disadvantage of the sub-dewpoint processes. It is an object of the present invention is to overcome this sensitivity towards the correct amount of combustion air.

During the switching of the sub-dewpoint reactors, a decrease in recovery efficiency is experienced for a 3-reactor plant (one conventional Claus reactor and 2 sub-dewpoint reactors)- This is caused by the fact, that the first step in the reactor switching process is the opening of the valve in the hot feed line to the cold (final) reactor which reactor at that time has a reduced conversion capability due to its pores being filled with liquid sulfur. Most of the gas will then begin to pass only partly converted through thh deactivated cold converter and from the cold condenser to the incinerator. At this point the plant operates effectively more like a 2-reactor MCRC unit then lite a 3-reactor unit and the sulfur recovery will be reduced.
The sub-dewpoiat processes are equipped with switching valves for the batch-wise regeneration of the sulfur loaded catalytic aub-de**rpoint reactors. The switching valves are sensitive for leaks. Liquid sulfur may combine with soot, salts, refractory dust, catalyst dust and pipe scaling to form sulfur concrete. This very hard sulfur concrete may precipitate snd/or collect on the valve seat, causing the valve not to close completely. Consequently, the valve will start to leak. Process gas will bypass the final sub-deVpoint reactor and the sulfur recovery wiB. drop.
The overall sulfur recovery efficiencies of the sub-dewpoint processes are normally in the range of 98.5-99.5%, depending on the acid gas feed composition and the number of catalytic (sub-dewpoint) reactors; Very high recoveries, in the range of 99.5-99,7%, cannot he reached. As the requirements
of the authorities will increase in the future to limit the amount of emitted SO2
there is a need for processes with an increased sulfur recovery efficiency.
In US patent specification no. 4,988,494, it is proposed that the H2S concentration, in the gas leaving the last catalytic Claus stage is tontrolled to have a value ranging between 0.8 and 3% by volume by reducing the quantity of combustion or oxidation air passed to the oxidation, stage. The H2S in the residual Claus gas is subsequently converted -with high selectivity to sulfur in a dry-bed oxidation stage.

The increase of the HaS concentration will result in a decreased SO2 concentration, however, not to very low levels. For an H2S concentration of 0.8% by volume, the SO2 concentration will be typically 0.03-0-15% by volume, and this will result in a sulfur recovery efficiency loss of typically 0.09-0.45%. As SO2 is not converted in a dry-bed oxidation stage, this will result in appreciable sulfur recovery losses, and consequently sulfur recovery efficiencies close to 100% cannot be reached.
A second disadvantage of operating with excess HaS compared to SO2 is that the temperature increase in the dry-bed oxidation reactor becomes higher with increasing H2S concentration.
Higher reactor temperatures will result in an increased formation of SO2 as a result of gas phase and catalytic osridation of formed sulfur vapor. Also for this reason, a shifted operation towards increased concentrations of HaS in the outlet of the last Claus converter is not beneficial.
It has been experienced, that if the catalyst bottom temperature in a dry-bed oxidation reactor exceeds 250-260°C, the H2S oxidation efficiency to elemental sulfur will start to drop from 94-96% to lower values. Combined with a reactor inlet temperature of approximately 1SO-2OO°C? this results in an upper limit for an acceptable temperature increase, due to the reaction heat developed, of some 60-80°C, corresponding with 0,8-1,1 voL% of SsS in the process gas to the dry-bed oxidation stage.
The shifted operated sulfur plant, followed by a dry-bed oxidation step with an oxidation catalyst which is not effective in promoting the Claus reaction, is known as the 3UPERCLAUS® or SUPERCLAUS®-99 process.
The SUPERCLAUS® process, as well as the SUPERCLAU3-99.5 process, is described in "SUPERCLAUS® * the answer to Claus plant limitations", Lagas, JA.; Borsboom, JM Berben, P.H,, 38th Canadian Chemical Engineering Conference, Edmonton, Canada.
Surprisingly it has now been found that it is possible tt> increase the sulfur recovery by very simple process modifications. According to the

invention a process for the recovery of sulfur from a hydrogen containing gas, comprises;
a. Oxidizing part of the hydrogen snlfide in a gaseous stream with
oxygen or an oxygen contairang gas in sin oxidation stage to sulftiz dioside, and
thereafter reacting the major part of the remaining hydrogen eulfide and the
major part of sulfur dioxide to elemental sulfur and water;
b. reacting the product gas of oxidation stage a) in at least three
catalytic reaction systems
wherein
- at least one system (i) is operating above the sulfur dewpoiat ih accordance
with the Claus equation
2 H2S + SO2 o 2 H2O + 3/n Sn;
- at least one system (ii) is operating at a sub-dewpoint temperature in
- accordance with the Claus equation, downstream of system (i); and
- at least one system (iii) is regenerating above the dewpoint of sulfur or
- operating in accordance with the Claus equation above the dewpoint of sulfur,
- usually down stream of system (i) or off-line; and
c> selectively obelizing HaS in the gas leaving stage b i in particular leaving the (last) sub-dewpoint Claus reactor) to elemental sulftar, preferably employing for this purpose a catalytic stage including a selective oxidation catalyst, which is substantially insensitive to the presence of water vapor in the gas stream and is ineffective in promoting the establishment of the Claus equilibrium: 2 H2S + SO2 c=> 2 H2O + 3/n Sn-
System (i) in stage b may very suitably be a Claus realtor known in the art ( i.e. a conventional reactor) operating under conditions known, in the art.

Usually, the reactor(s) ia stage b) that operate sub-depoint periodically switch between the eub-dewpoint operation (as system ii) and regeneration (as system iii),
The regeneration of system (iii) is typically carried otft to vaporize the liquidsulfur collected in the reactor during sub-dewpoint operation. After removal of the vaporized sulfur, the system (iii) may switch to step-dewpoint operation and serve as system (ii).
System (iii) may be regenerated off-line or on-line, preferably online, e.g. as indicated below in the description and the figures, ift particular in an on-line configuration, system (iii) is usually situated upstream of system
(ii).
A process according to the invention may very suitably be carried
out without removal of water prior to stage c).
It has been found that in accordance with the invention it is possible to reduce or even nulliiy the decrease in recovery efficiency due to the switching of the sub-dewpoint beds from the Glaus conversion stage to the regeneration stage and visa versa, as for instance in a conventional MCEC process with one catalyst bed in the sub-dewpoint mode.
It has further been found that a process according to the invention mitigates the sulfur recovery decrease, which would otherwise otcur in a conventional sub-dewpoint process due to equipment problems such as leaking switching valves
In addition the present invention allows the increase of the sulftir recovery efficiency of the sub-dewpoint process compared to a known sub-dewpoint process.
It is preferred in said step c) of selectively (radioing H>J3 also employing a stoxchiometric excess of oxygen sufficient to result iti an overall excess of oxygen being employed in the total process for the recovery of sxilfur from the hydrogen sulfide containing gas.

It should be noted, that complete conversion/removal of SO %> ia the residual Claus gas, down to the ppmv level ae in the SCOT process, is not required in case the sub-dewpoint Claus reactor step is followed by a dry-bed oxidation step. Conversion of SO2 to sulfur vapor down, to a level of approximately 100 ppmv (typically to about 20-200 ppmv) is acceptable and will result in negligible recovery efficiency losses.
In the SUPERCLAUS® process, as described in US-A 4,988,494, a significant residual concentration of SO2 is left in the process^^to the dry-bed oxidation stage. This residual SO2 not only decreases the overall sulfur recovery efficiency, because it is not converted to elemental sulfiar, but also decreases the activity of the selective oxidation catalyst.
If one wants to overcome this decrease of activity in the known conventional SUPERCLAUS® process, the temperature level in the dry-bed oxidation reactor is preferably increased, but a higher temperature level will decrease the oxidation efficiency to elemental sulfur. Consequently, very high oxidation efficiencies in the dry-bed oxidation stage, in the ranga of 94-96%, are not possible with process gas containing concentrations of SOz in the range of 300-2000 ppmv.
The control of the process with an incorporated low tefriperatuxe sub-dewpoint Claus reactor is also much more fle^dble. A varying SO* content in the process gas from the last (conventional) catalytic Claus reactor (system (i), operating above dewpoint) will not result in varying recovery losses by SO2 slippage, because basically all SO2 can be converted to elemental sulfur in the sub-dewpoint stage* This makes the control on H2S much lees sensitive to process fluctuations. As in the SUPERCLAUS process, the ratio H2S to SO2 is more than 2- This excess of HsS suppresses the SO2 content to v6ry low levels.
In a process according to the present invention, the concentration of the hydrogen sulfide gas in the residual gas can be controlled in & surprisingly simple way. Thus, for example, the signal from an H2S analyzer £n the residual

gas can be used to set or adjust the amount of combustion air to the Claus burner, or oxidation air to the dry-bed oxidation stage, such that the process gas leaving the sub-dewpoint stage is in excess, ie- the HaS/BOg. ratio is larger than 2. Such angles would normally be considered unsuitable for a
sub-dewpoint dans process because a sub-dewpoint Claus process operates, like the conventional Claus process, at its optimum if the ratio H2S/SO2 in the tail gas equals 2, regardless of the temperature of the final reactor. In contrast, the optimum EfeS TOn^ntration for the SUPERCIAUS feed gas is dependent on, among others, the temperature of the preceding Glaus stage,
and is in a normal SUPERCLAUS plant kept at a fixed value because the temperature of the preceding Claus stage is virtually constant. During switching of the reactors in the process of the current invention however, the normal SUPERCLAUS control philosophy of a fixed H2S concentration in the Claus tail gas would result m a recovery less than that of the sub-dewpoint process alone. This would be a serious disadvantage of the present inv ention. It is an object of the present invention to minimize the recovery losses during switching, byj^exagorarily adjusting the H2S setpoint to a larger value, 0.5 -3.0 vol% preferably 1 -1.5 vol%, as long as the switching process continues, ie. as long as the switching valves are being opened and closed, respectively. Normally in a conventional SUPERCLAUS plant, the H2S setpoint is preferably kept below 1.0 vol% to prevent excessive SOa formation at the resulting high bottom temperatures in the selective oxidation stAge. However, it has been found that a higher H2S setpoint during a relatively short time such as required for switching the sub-dewpoint reactors does net lead to excessive SO2 formation.
It has been measured in aja MCRC unit, processing requiem acid gas, that the SO2 content of the process gas at the outlet of the last step-dewpoint Claus reactor, containing an H2S concentration of 0.50 vol.% (wet basis), was 11 ppmv. This extremely low level of SO2 is caused by the excess HaS operation and the low reactor temperature, which suppresses the SO2 content further,

i.e, the Glaus equilibrium, of equation (2a) is almost completely on the sulfur side. This very low SO2 level results in a negligible sulfur recovery efficiency loss. Also, a fluctuation of this H2S content, say 0.5 ± 0.2 %, will have a marginal effect on the SO2 content and consequently on the SO* sulfur recovery efficiency loss, This effect stabilises the upstream part (thermal stage, first Glaus reactor stage (system i), sub-dewpoint reactor staged) of the sulfur recovery unit, i.e. a fluctuating excess H2S in the MCRC tail gara will have only a marginal (very small) effect on the SO2 content and the 8O2 sulfur recovery efficiency losses. This excess H2S also mates the control of the Combustion air much more insensitive with respect to the plant performance recovery efficiency*
A process according to the invention can be suitably applied for the treatment of gases containing hydrogen sulfide, but also for gasses containing both hydrogen sulfide and substantial quantities of ammonia (c£ NL-C-176160), in the latter case, the temperature in the combustion chamber is preferably at least 1250°C.
Under stoic metric combustion in the Claus burner generates excess H2S over SO* (molar ratio > 2:1). A disadvantage of operating with excess H2S over SO2 in the Claus tail gas to reduce the SO2 content, as in the SUPERCLAUS® process, is that this operation mode results in less combustion aix to the main (H2S) burner compared to the conventional model of operation with H?J3;SO2 - 2:1. This will result in a decrease of temperature of combustion in the combustion chamber, which is detrimental to the destruction efficiency of ammonia, which requires high temperatures. Sub- dewpoint Claus conversion, before dry-bed oxidation, allows for more o:xygen to *;he Claus burner, and therefore for higher combustion temperatures* white maintaining very low SO2 levels in the residual Claus gas to the dry bed oxidation stage.
In stage a) (the so called thermal stage) part of the hydrogen sulfide is preferably oxidized to sulfur dioxide, under the conditions thdit are comparable to a conventional SUPERCLAUS or sub-dewpoint process.

In stage b) the sub-dewpoint reactor or reactors is/are preferably operated under the following conditions.
The inlet temperature is preferably controlled and is preferably at least about 120 °C , more preferably at least 125 °C.
The outlet temperature is resulting and is typically about 150 °C or less.
Very good results have been achieved wherein the sub-dewpoint reactor comprises Glaus catalysts with an increased macropore Volume and an increased total pore volume.
In stage b) preferably at least one reactor is operating as a Clans reactor under sub-dewpoint conditions while at least one other reactor is being regenerated. The reactor operating under sub-dewpoint conditions and the regenerating reactor preferably comprise the same catalytic material After an appropriate period the reactors preferably are switched, such that the regenerated reactor takes over the Glaus reaction and the other Reactor is regenerated.
In a process according to the invention, the hydrogen sulfide gas remaining in the residual gas can be processed after stage b to form sulfur by a known per ee method. Such methods are described in the literature. Preferably, however, the remaining gaseous hydrogen suMde ie oxidized with air in an oxidation stage to form elemental sulfur in accordance with the following reaction:
It has been found that, when the concentration of the hydrogen sulfide leaving the last catalytic sub-dewpoint Claus stage is maintained at a value of between 0.1 and 0.5% by volume optimum sulfur recovery percentage, of 99.5-99.8 can be obtained, after selective oxidation, The oxidation can suitably take place by dry-bed oxidation or by oxidation in a

liquid, in which, in general, sulfur and water vapor have first faeen removed from the residual gas,
In particular in case dry-bed oxidation is used in stage c, the HsS concentration in the residual gas is preferably maintained between 0.2 and 0.4% by volume, because up to about 0.4% by volume of HaS a tery satisfactory total sulfur recovery percentage is observed.
In a dry oxidation bed, the oxidation to sulfur can be effected by a known per se method using an oxidation catalyst. One example of an oxidation catalyst and the application thereof is described in US-A-4311683,
The method described therein is the Selected process (Hass, R.H.; Ingalis, M.N.; Trinker, T.A.; Goar, B.G., Purgason, R.S.S,: "Process meets sulfur recovery needs", Hydrocarbon Processing, May 1981, pages 104-107). In this process, HsS is oxidized to S and SO2 using a special catalyst such as vanadium pentoxide on. alumina. Approximately 80% of the HaS supplied is oxidized to elemental sulfur, if water vapor is removed to a substantial extent. Another application of a dry-bed process which is not sensitive to water vapor in the process gas is the absorption of H2S in an absorption mass as described, for example, in European patent no- 71983.
Another type of catalyst, which may be applied for that dry-bed oxidation, comprises a carrier of which under the reaction conditions applied, the surface exposed to the gaseous phase does not exhibit activity for the Claus reaction.
Such a catalyst may be one described in US patent specification number 4,818,740 and 5,286,697 or in WO-A 9732813.
Very good results have been achieved with a catalyst for the selective oxidation step c, comprising a carrier material of which under the reaction conditions applied, the surface exposed to the gaseous phase does not exhibit activity for the Claus reaction* Preferred examples of such carriers include, inter alia, silica and alpha-alumina

The catalytically active material of the oxidation catalyst is preferably present on the carrier in a proportion of 3-10% by weight calculated on the total mass of the catalyst. Suitable examples of catalytically active materials include metal oxides. The metal oxide may be an oxide of only one metal, a mixed oxide of two or more metals or a mixture of metal oxides. Very suitable is a metal oxide comprising iron oxide. Very good results have been achieved with a mixed oxide of iron and at least one second metal, for instance chromium or zinc.
The BET specific area (e.g, as described in E* Robens et al.: "Standardization of sorption measurements and reference materials for dispersed and porous solids", Ch- 3 o£ A. Dabrowski (ed-): "Adsorption and its Application in Industry and Environmental Protection". Vol. 1: ^Application in Industry"; Studies in surface science and catalysis, Vol. 120A; SJlsevier, Amsterdam 1999) of the catalyst is preferably more than 20 ma/tj catalyst. The upper limit is not particularly critical. Very good results have b«^en achieved with a BET specific area in the range of about 30 to about 120 xA2/g.
The average pore radius is preferably at least 25 A, BA denatured by mercury intrusion porosimetry, more preferably about 100 - 700 A.
It has been found that - with a view to a maximum sulfur recovery percentage - the choice of the optimum volume percent of HaS ii* the residual gas is depending on the extent of the efficiency of the last oxidation from HsS to sulfur ia the dry oxidation bed.
The minimum volume percent of H2S, corresponding with the maximum volume percent of SO& in the residual gas from the ldst Claus stage, is determined by the amount of SO2 in the residual gas. This should be low enough, in practice typically lower than 200 pptav, in order to result in very small sulfur recovery efficiency losses. The corresponding minimum volume percent of H2S is then usually approximately 0.2 voL%
For this reason (the upper SO2 limit), the H2S volume percent is generally preferred not to be lowered too much. Also, when the &2S volume

percentage to the dry-bed oxidation reactor is too low, the oxidation efficiency to sulfur in the dry-bed oxidation reactor is not at its
When the HsS volume percentage is'too high, the overall recovery efficiency will also decrease. For above reasons, the H2S volume percentage in the process gas from the last Claus stage, is preferably in the range of about 0.2-0.4 voL%.
Liquid oxidation, too, can take place using a known process. Examples of known processes are the Stretfbrd process (the Chemical Engineer* February 1984, pages 84 &), the Lo-Cat process of Aifi Resources Inc. or the Takahax process.
The control of the oxidation air to the selective oxidation stage is not critical and can thus be kept simple.
A process according to the invention can be carried otlt in an existing sub-dewpoint Claus plant and requires only relatively simple modifications of the existing control of the gas streams. In case a 3-stage sub-dewpoint Claus plant (with two switching sub-dewpoint reactors) is used, a selective oxidation reactor is provided in the specific embodiment of the present indention, which in relation to the cost involved in other residual gas processing plants is inexpensive. Thus the application of the process according to thft present invention leads to considerable economic advantages.
In case a 4-stage sub-dewpoint Clans plant is used only the fourth catalytic reactor needs to be arranged as a selective oxidation reactor. In this case, too therefore, a considerable economic benefit is obtained.
An additional benefit is that the negative effect of the switching of the sub-dewpoint reactors on the H2S and SO2 content in the tail gas of the sub-dewpoint part, which has the tendency to increase at this switching, resulting in a drop in recovery efficiency, is suppressed by the operation with excess HgS- The SO2 level will increase, but only marginally, without affecting the recovery losses caused by this SO 2. In the meantime the Kz& content will not increase because it is on analyzer setpoint control.

Another beneficial effect is that the sub-dewpoint part is operated on

excess H2S mode. This will result in

a more reducing process gas, which is very

beneficial for the Claus catalyst activity. In this way the sulfattf content of the Claus catalyst is maintained at a very low level, resulting in a higher catalyst activity.

condensed from the gas flows using conventional systems, such as condensers. In case an especially high sulfur recovery is neceBsaiy it may be advantageous to use the system disclosed in EP-A 655,414, more in particular for the
treatment of the final product gas
-, after the last treatment, A beneficial effect is that the sulfur production load stufts more to the last reactor stage, Le. the selective oxidation reactor or SUP&RCLAUS® reactor. This means that the sulfur production of the sub-dewpoint reactor(s) will decrease from approximately 9% of the sulfur quantity in tHe plant feed, to approximately 7% of the sulfur plant load. Accordingly, the sub-iewpoint reactor can be operated for a longer period of time before it has to be regenerated. Alternatively, less catalyst can be applied in this reactor for the same absorption period.
in the selective oxidation stage with overall SEE of typically 99.5-99.8%.
A process according to the present invention will now be described in more detail with reference to the accotaxpanjdng figures 1, 2 and Ji AH parameters (such as temperature) are given by means of preferred examples.
As shown in figure 1, the feedstock gas (=Claus gas) is Supplied through line 1 to the Claus burner/combustion chamber 2. The ar&ount of
The most important beneficial effect is, however, that the sulfur recovery efficiency will increase significantly- The sub-dewpoint configuration of MCRC/CBA/Sulfreen/Maxisulf .removes the bulk of the feed stdfiSr and produces a process gas with basically no SO2 (ppmv level) at a relatively low amount of H2S, typically 0.2-0.4 vol.%. This H2S is oxidized to elemental sulfur
high efficiency, resulting i*i a high

combustion air, controlled by the qulantity proportion regulator 3 and H2S analyzer 28, is supplied to Claus buicner/combustion chamber 2 through line 4. The heat generated during the combustion (1200°C) of the Class gas is dissipated ia a boiler 5 heated by sp^nt gases, producing steattr that is discharged through line 6.
The Claus reaction takes place in the burner aixd the combustion chamber. The sulfur formed is condensed in boiler 5 (150°C) and discharged through line 7. The gas is passed through line 8 to a heater 9 where it is heated to the desired reaction temperature of 250°C before beitfg supplied through line 10 to the first Glaus reactor 11. In reactor 11 the Claus reaction

takes place again, whereby sulfur is

formed. The gas ie discharged through

line 12 to the sulfur condenser 13. The condensed sulfur (150°d) is discharged through line 14. Thereafter the gas is passed through line 15 to the next

reactor stage, which again includes *

heater 16, a reactor 17 and a sulfur

condenser 18, In this reactor the Claus reaction takes place again. In the MCRC process configuration, Claus reactor 17 is in the regeneration mode> at an elevated temperature of 300-350°C, to vaporise the liquid sulfur collected in the Claus catalyst.
The condensed sulfur (125^0) is discharged through line 19. The steam generated in the sulfur condenser is discharged through lines 20, 21 and 27.
The process gas is passed through line 22 to the last Olaus reactor

23. This reactor is operating at a low

temperature (sub-dewpoint mode). Note

that for the MCRC process configuration no heater upstream of this reactor is required. The sulfur is condensed in iulfiir condenser 24 (125°C), and discharged through line 25.
The H2S concentration in the residual gas line 26 is cctotrolled by an

H2S analyzer 28 to a range from 0.1-*

.5 % by volume. The H^S analyzer

controls a control valve in combustioxlL air line 29.

The residual gas is passed through line 26 to the sulphur removing stage 3 The gas is then passed though line 33 to an after-bu-tner 34 before being discharged through chimney 35.
As slaown in figure 2, a lean Claus feedstock gas is supplied through line 1 to an oxidation reactor 2. An amount of oxidation air controlled by the quantity ratio regulator 3 and H2S analyzer 24 is passed to the foddatioa reactor through, line 4* In the oxidation reactor a portion of the H2S is oxidized over a special catalyst to form SO2 wtiere after the Claus reaction takes place. To prevent an unduly high temperature from arising within the oxidation reactor, as a result of the reaction heat, a quantity of gas is recycled by means of blower 6 through lines^ 5 and 7. The gas from reactor 2 is passed through line 8 to sulfur condenser 9, where tie sulfur formed during the reaction is condensed at 150°C and discharged through line 10. The heat generated during the reaction is dissipated in sulfur condenser 9 with generation of steam, which is discharged through hue 11. The gas is passed through litxe 12 to heater 13, where it is heated, for example, to 300DC before supplied to Claus reactor 14, which is being regenerated from the collected sulfur. In condenser 15, the sulfur is condensed at 125°C and discharged through lino 16, and steam generated is discharged through line 17.
The process gas is passed through line 18 to the last Claus reactor 19. This reactor is operating at a low temperature (sub-dewpoini mode).
In condenser 20, the sulfur is condensed at 125*C, and discharged through line 21. Steam is discharged through line 22.
.5 percent by volume. The fckS analyzer
The H2S concentration in tne residual gas line 23 is controlled by an H2S analyzer 24 to a range from 0.1-C
controls a control valve in the combustion, air line 25.

line 29 to an affcer-trumer SO before Figure 3 shows in greater bed and the oxidation in a liqirid as
The residual gas is passed through line 23 to the sulfur removing stage 26. The air required for the oxidation is supplied through line 27. The sulfur formed is discharged through! line 28. The gas is then parsed through
being discharged through ckimney 31.
detail the oxidation or absorption in a dry-indicated more generally it* 30 of figure 1
or 26 of figure 2.
In figures 1 and 2, the residual gas is supplied through lines 26 and 23, respectively, in figure 3, the residual gas is supplied through line 1.
In figure 3a, after the removal of the sulfur from the Residual gas in separator 2, which is discharged through line 3, and the condensation, of the water in 4, which is discharged through line 5, the gas is supplied through a heater 6 to a selective oxidation reactor 7. The removal of sulfiif and water in 2 and 4 respectively, can take place using a known method, for example, as disclosed in US patent 4526590. In the selective oxidation reactor 7, a catalyst may be provided, for example, as described in the French paten* publications 8009126, 3105029 or 8301426. The required oxidation air ie supplied through line 8.
From the reactor, the gas flows to a sulfur condenser 9. The sulfur condensed is discharged through line 1O? and the steam generated through line 11. The
after-burner as designated by 34 in figiore
gas next flows through line 12 to the 1 and 30 in figure 2.
residual gas is supplied tfajftragh Hne 1
As shown in figure 3b, the
and heater 2 direct to the selective oxidation stage, that is to saf, without a preceding sulfur and water removing stage. This embodiment csto be used when a catalyst is present in the oxidation reactor 4, as described above, consisting of a non-Claus active carrier to which a,t least 0.1% bf weight of a catalytically active material, in particular a metal, oxide, has be6n applied, so that the specific area of the catalyst is more than. 20 m2/g, while the average pore radius is at least 25 A- The oxidation air required is supplied through line

3. The sulfcir condensed in sulftir condenser 5 is discharged though liae 6 and the steam generated through line 7J The gas next flows through line 8 to the after-burner designated by 34 in figure 1 and 30 in figure 2.

As shown in figure 3c the

residual gas is passed through line 1 to a

reactor 2 filled with an absorption mass, for example, as described in European patent no. 71983, published December 10+h, 1986. In reactor 2, the hydrogen

stdfi.de is removed from the residua!

gas by absorption- The gas next flows

through line 3 to the after-burner, designated by 34 in figure 1 and by 30 in, figure 2. When the bed is saturated, it is regenerated. Reactor 4 is connected in parallel to reactor 2 and is regenerated. By means of a circulation blowet 5, a quantity of gas is circulated- This gas is heated in heater 6, The air required for the oxidation is supplied througa line 7. The gas flows from Reactor 4 to sulfur condenser 8. The sulfur condensed is discharged through line 9 and the

steam generated through line 10. Tc

keep the system at the required pressure

a small gas stream is discharged through line 11 and recycled tft the feedstock

for the Claus plant (line 1 in figures

1 and 2).

As shown in figure 3d, sulfur is removed in separator 2, which is
discharged through line 3. Subsequently, in condenser 4, water is condensed which is removed through line 5- Th'* gas is passed to the liquid oxidation stage 6. The oxidation stage may contain, for example, a basic solution of sodium carbonate, ADA (anthraquinone disulfonic acid) and sodium metavanadate, as used in the well-known Stretford process.
KzS is absorbed in the liqiid and subsequently oxidized with air. The oxidation air is supplied through line 7 and the sulfur formed is

discharged through line 8- The gas n

flows through line 9 to **he after-

burner (34 in figure 1 ajid 30 in figure 2),
The invention, is illustrated by the following examples.

EXAMPLE 1 (FOR COMPARISON)
Using the apparatus as described with reference to figure 1, however, excluding the^ sulfurj^moval stage 30, the Glaus reaction is performed in a sulfur recovery unit having one catalytic Glaus siage operating above dewpoint and two -switching- reactors, of which the first seactor is in regeneration mode (above the dewpoint) and the other is in sub-ftewpoint mode. The Glaus reactors are filled with usual Claus catalyst su 4 vol.% CO2» 5 vol.% H2O and 1 voL% CSHG and 48.5 fctaoles/h O2 (a "deficit" of 0%) as air oxygen.
The tail gas composition after the third catalytic stag^ is as shown below. A total sulfur recovery efficiency (SRE) of 98.90% is obtained.


EXAMPLE 2 (FOR COMPARISON)
The Claus reaction is now performed in a sulfur recovery unit having one catalytic Claus stage operating above dewpoint and three -switching- sub-dewpoint reactors, of which the first reactor is in regeneration mode and the others axe in the sub-dewpoint mode (Other conditions as in i Example 1). Supplied to the thermal stage are a Claus gas, containing 90 voL% HaS, corresponding to 90 kmoles/h,
4 vol.% CO2, 5 vol.% H*O and 1 vol.% CsHs and 48.5 kmoles/h Oa (a "deficit" of 0%) as air oxygen.
The tail gas composition after the fourth catalytic stage is as shown below- A total sulfur recovery efficiency of 99.35% is obtained.


EXAMPLE 3 (FOR COMPARISON))
Using the apparatus as described with reference to figures 1 and 3b, the Glaus reaction is performed io. a sulfur recovery unit having two conventional catalytic Clans stages operating above dewpoint CtWet: first reactor 240 °C; second reactor 210 °C). Supplied to the thermal stage are a Glaus gas containing 90 voL% H2S corresponding to 90 kmoles/Ji, 4 vol.% CO2, 5 voL% H*O and 1 vol.% CsHe and 47.45 kmoles/h of O2 (a Mefidt" of 2.2%) as air oxygen- The H2S volume percentage in the residual gas afte* the second catalytic stage is 0.90.
The dry-bed oxidation is carried out using a water insensitive oxidation catalyst, comprising a silica carrier impregnated with iron oxide such as normally used in the SUPERCLAUS process. The inlet temperature is 210 °C. Using this catalyst with an oxidation efficiency of 86%, a total sulfur recovery percentage of 99*10% is obtained, corresponding with & tail gas composition at the outlet of the dry-bed oxidation stage as showfa below.


EXAMPLE 4 (FOR COMPARISON)
The sulfur recovery unit is now equipped with three Conventional catalytic Claus stages operating above dewpoint and one oxidation stage. The inlet temperatures of the reactors are as in the previous example, the third reactor has an inlet temperature of 195 °C. The amount of air O2 is 47.55 kmoles/h (a "deficit" of 2.0%). The H^S volume percentage in th# residual gas after the third catalytic stage is 0.70.
The dry-bed oxidation is carried out using a water insensitive oxidation catalyst with, an oxidation efficiency of 86%, A total sttffur recovery percentage of 99.40% is obtained, corresponding with a tail gas composition at the outlet of the dry-bed oxidation stage as shown, below.


Using the apparatus as described with, reference to figure 1 and 3b, the Glaus reaction is performed in a sulfur recovery unit having one catalytic Claus stage
operating above dewpoixtt (e.g. about 240 °C) and two -switching- sub-dewpoint reactors, of which the first reactor is in. the regeneration mode ^at about 300 °C) and the other is in the sub-dewpoint mode (at about 125 °0). Supplied to the thermal stage are a Glaus gas, containing 90 voL% HaS, corresponding to 90 kmoles/h, 4 vol.% CO*, 5 voL% H2O and
1 vol-% C2H6 and 43.S moles/h O2 (a "deficit" of 1 %) as air oxygen. The H*S volume percentage in the residual gas after the third catalytic stage is 0.30, the SO2 volume percentage is 0.0025.
The dry-bed oxidation is carried out (at 210 *C at hxkt) using a water insensitive oxidation catalyst with an oxidation efficiency of 86%. A total sulfur recovery percentage of 99.66% is obtained, corresponding with a tail gas composition at the outlet of the dry-bed oxidation stage as shovm below.








New Claim
1. Process for the recovery of sulfur from a hydrogen sulftde containing gas, comprising:
a. Oxidisring part of the hydrogen sulfide in a gaseous stream with
oxygen or an oxygen containing gas in an bxidation stage to sulfur dioxide, and
thereafter reacting the major part of the remaining hydrogen sulfide and the
major part of sulfur dioxide to elemental stdfur and water,
b. reacting the product gas of oxidation stage a) in at feast three
catalytic reaction systems
wherein
- at least one system (i) is operating above the sulfur deWpoint in
accordance with the Clan© equation
2 H2S + SO2 2 H2O + 3/n Sn;
- at least one system (ii) is operating at a sub-dewpoint temperature in
- accordance with the Claus equation, downstream of systein (i); and
- at least one system (iii) is regenerating above the dewpoint of sulfur or
- operating in accordance with the Claus equation above thft dewpoint of
- sulfur,
and wherein the amount of oxygen or oxygeti containing gas of stage a) is adjusted, such that the process gas leaving the sub-dewpoint stage has an HaS/SOg ratio of more than 2;
c. selectively oxidizing H2S in the gas leaving the (last) sub-dewjppint Claus
reactor of stage b (in particular leavtag) tho-^asfr) eiab dewpemt ■Ck^.
to elemental sulfur, preferably employing fof this purpose a. catalytic stage including a selective oxidation catalyst, which is substantially insensitive to

the ^bhehment of the Claus equilibrium: 2 H,S + SO2 « 2 H.0 + 3/n a,
2. A process as claimed in claim 1, wherein the selective oxidation is effected in a
3. dry oxidation bed,
4. A process as claimed in any of claims 1 or 2, wherein a catalyst is used for the
5. selective oxidation step la, which catalyst preferably comprises a

carrier material of which, under the reaction conditions applied, the surface
exposed to the gaseous phase does not exhibit activity Jbr the Clans reaction,
and a catalytically active material, the specific area of the catalyst being more
than 20 m*/g catalyst, and having am average pore radius of a* least 25 A*
4- A process as claimed in claim 3, wherein the gas obtained in step b has
an H2S concentration of 0.1-0,8 % by volume and the osidatioi* efficiency to
siiifur of the oxidation catalyst of 80-96%.
5. A process according to claim 4, wherein the H2S concentration is 0.1-0.5,
preferably 0.2-0.4 % by volume.
6 A process as claimed in claim 3-5, wherein the catalyst contains silica or
alpha-alumina as carrier material*
7. A process as claimed in claim 3-6, wherein the catalytically active
8. material of the oxidation catalyst is present on the carrier in a proportion of 3-
9. 10% by weight calculated on the total mass of the catalyst.
10. A process as claimed in any of the claims 3-7, wherein tile catalytically
11. active material is a metal oxide.
12. A process according to claim 8, wherein the metal oxide is a mixed oxide
13. of two or more metals, or a mixture of metal oxides.
14. A process as claimed in claim 8 or 9, wherein the oxide comprises iron
15. oxide
16. A process according to claim 10, wherein the oxide is a mixed oxide of
17. iron and at least one element selected from the group consisting of chromium
18. and zinc.
19. A process according any of the preceding claims, where the H2S
20. concentration of the gas obtained in step b is adjusted during the switching of
21. tlie sub-dewpoint reactor(s).
22. A process according claim 12 in which said HsS concentration is kept
23. between 0.5 and 3 vol%, preferably between 1.0 and 1.5 vol%.

24. A process according to any of the preceding claims, wherein in systems
25. (ii) and (iii) reactors are used that alternatingly operate sub-deitfpoint and
26. regenerate at a temperature above the dewpoint.
27. A process according to any of the preceding claims, wherein system (iii)
28. is regenerated in-line with the gas stream.
29. A process according to any of the claims 1-14, wherein system (iii) is
30. regenerated off-line, preferably in a closed loop.
31. Sulfur recovery installation for carrying out a method according to any
32. one of the preceding claims, comprising at least one hydrogen stilfide oxidation
33. unit; at least one Claus reactor designed to operate above the stttfiir dewpoint,
34. at least two Clans reactors designed to be alternatingly operating sub-
35. dewpoint and regenerating above dewpoint; and downstream of said Claus
36. reactors at least one unit for selectively osidizing H2S.
4#


Documents:

2563-CHENP-2006 AMENDED PAGES OF SPECIFICATION 21-06-2011.pdf

2563-CHENP-2006 AMENDED CLAIMS 21-06-2011.pdf

2563-CHENP-2006 EXAMINATION REPORT REPLY RECEIVED 21-06-2011.pdf

2563-chenp-2006 form-1 21-06-2011.pdf

2563-chenp-2006 form-3 21-06-2011.pdf

2563-CHENP-2006 CORRESPONDENCE OTHERS 05-10-2010.pdf

2563-CHENP-2006 CORRESPONDENCE PO.pdf

2563-CHENP-2006 FORM-18.pdf

2563-CHENP-2006 POWER OF ATTORNEY 05-10-2010.pdf

2563-chenp-2006-abstract.pdf

2563-chenp-2006-claims.pdf

2563-chenp-2006-correspondnece-others.pdf

2563-chenp-2006-description(complete).pdf

2563-chenp-2006-drawings.pdf

2563-chenp-2006-form 1.pdf

2563-chenp-2006-form 3.pdf

2563-chenp-2006-form 5.pdf

2563-chenp-2006-pct.pdf


Patent Number 248726
Indian Patent Application Number 2563/CHENP/2006
PG Journal Number 33/2011
Publication Date 19-Aug-2011
Grant Date 11-Aug-2011
Date of Filing 13-Jul-2006
Name of Patentee JACOBS NEDERLAND B.V.
Applicant Address Plesmanlaan 100, NL-2332 CB Leiden
Inventors:
# Inventor's Name Inventor's Address
1 BORSBOOM, Johannes Frans Halskade 104, NL-2282 VC Rijswijk
2 VAN WARNERS, Anne Cornelis Vlotstraat 13, NL-1318 AE Almere
3 VAN NISSELROOIJ, Petrus, Franciscus, Maria, Theresia Gerard Noodtstraat 115, NL-6511 SV Nijmegen
4 VAN YPEREN, Renee Zonnedauw 7, NL-6961 PL Eerbeek
5 CHOPRA, Vijay, Kumar 98 Falworth Way NE, Calgary, A T3J 1Y1
PCT International Classification Number C01B17/04
PCT International Application Number PCT/NL2005/000023
PCT International Filing date 2005-01-14
PCT Conventions:
# PCT Application Number Date of Convention Priority Country
1 04075104.2 2004-01-16 EUROPEAN UNION