REFINERS FOCUS ON FCC, HYDROPROCESSING, AND ALKYLATION CATALYSTS

March 28, 1994
Refiners and technical specialists from around the world met last fall in Dallas at the National Petroleum Refiners Association Q&A (question and answer) meeting to exchange experiences and discuss problems in key operating areas. At this renowned annual meeting, held Oct. 20-22 in 1993, a panel (see box) of industry experts responded to presubmitted technical questions. Additional questions and comments were then invited from the floor by moderator and NPRA technical director Terrence Higgins.

Refiners and technical specialists from around the world met last fall in Dallas at the National Petroleum Refiners Association Q&A (question and answer) meeting to exchange experiences and discuss problems in key operating areas.

At this renowned annual meeting, held Oct. 20-22 in 1993, a panel (see box) of industry experts responded to presubmitted technical questions.

Additional questions and comments were then invited from the floor by moderator and NPRA technical director Terrence Higgins.

This first installment of excerpts from the conference transcripts includes exchanges on:

  • Effects of fluid catalytic cracking (FCC) catalysts on light olefins yields.

  • Experiences with removing sulfur contamination from steam methane reforming catalyst by in situ regeneration.

  • Selection of appropriate uses for regenerated hydro-desulfurization catalyst.

  • Increases in reformate and hydrogen yields using bimetallic catalyst at reduced pressure.

  • Effects of sulfate formation during reforming catalyst regeneration.

  • Solid alkylation catalysts.

LIGHT OLEFINS YIELD

Please discuss various options for increasing light olefin yields, including: ZSM 5; USY or other higher silica alumina ratio catalysts; other additives; ultra-short contact time; etc. What have been the maximum light olefin and isobutylene yields observed from the fluid catalytic cracking unit (FCCU)?

Bonelli: At one of our refineries, in an effort to increase isobutylene yields without violating an olefin processing constraint, we have tried several catalyst combinations and additive combinations. The unit in question has a minimum riser time of about 1.5 sec. We tried high rare earth catalysts in which our base LPG yields were 23 24 liquid vol % and isobutylene was 22.1.

We did try the ZSM 5 additive and wound up with LPG yields increasing to about 30 31 vol % liquid, and isobutylene up to about 3. We have since changed catalyst, gone to a lower-rare earth catalyst and are currently operating with LPG yields in the 26 27 LV % range with isobutylene of about 2.8. So the ratio is favorable.

Juno: Several years ago we did a test run adding ZSM 5 to our catalyst inventory. We experienced higher FCC gasoline octane, more light ends, greater olefin production, but less FCC gasoline with a very waxy feedstock.

O'Brien: A fresh catalyst composed of USY zeolite vs. REY zeolite will yield a product with higher olefins. The USY zeolite produces more olefins due to reduced hydrogen transfer reactions. The increase in olefin yields will be unit specific, depending on feed type, operating temperature, and residence time in the riser and reactor.

We tested ZSM 5 in our pilot plant units and the trends were as expected. Gasoline yield was decreased with increased yield of alkylation feedstock. Gasoline components reduced in the ZSM 5 product are n-C7 through n C11 olefins and paraffins, The maximum yield of butene 1 plus isobutylene in a Phillips Petroleum Co. heavy oil cracker (HOC) is 5 wt % of fresh feed. Approximately one third of this is isobutylene.

Sloan: Options for increasing, light olefin yield fall into categories of catalyst, process operating conditions, and hardware. Operating variables and catalyst play a decisive role in determining how much light olefin can be formed in the first place.

Process equipment design, in addition to catalyst, is important in determining how much olefin yield is retained against secondary hydrogen transfer reactions. All of these factors are interrelated.

Catalyst options include:

1.Controlling catalyst unit cell size and rare earth level to minimize hydrogen transfer. (Ultrastable Y zeolite catalyst is preferred. This is a complex issue in itself and really goes beyond the scope of this question.)

2.Using ZSM 5 additive, which converts low octane olefins in the gasoline boiling range to light olefins.

The primary operating variable affecting light olefin formation is reactor temperature. As cracking temperature is increased to increase conversion, light olefin yield increases. Beyond the peak gasoline yield, it increases very rapidly with conversion.

Design variables affecting light olefin yield include reaction time and pressure.

1.Reaction time Catalyst and oil vapor contact time should be short to minimize hydrogen transfer. Particularly, long residence time, low velocity, back mixed conditions in the disengager dilute phase need to be avoided. Also, short contact time in the riser involves more than just restricting the physical volume of the riser. Design of the catalyst delivery and feed injection systems should be optimized to prevent zones of high catalyst slippage and refluxing within the riser itself.

2.Pressure The independent effect of hydrocarbon partial pressure is higher C3 and C4 olefinicity with lower partial pressure. Partial pressure can be lowered with additional dispersion steam usage, which may be a design variable or an operating variable, or through operating at low pressure. Since most operating units do not have the flexibility to reduce pressure, this needs to be considered during the design of the unit.

It is also worth pointing out that the higher cracking temperatures mentioned earlier require a reaction system which will minimize dry gas formation. This means avoiding thermal cracking at the feed injection point and in the disengager. A state-of the art atomizing feed injection system is needed in the mix zone of the riser, and closed cyclones are needed in the disengager to fulfill this requirement.

A paper presented at this year's NPRA annual meeting gave some estimates as to what modern hardware and catalyst technology could do in a unit purposely designed to maximize light olefins.

The results showed that with 20% ZSM 5 in the catalyst inventory, and using high temperature, short contact time cracking, potential MTBE yield could be 6.6 vol % and potential alkylate yield could be 61.9 vol % of FCC feed. These figures corresponded to an isobutylene yield of 5.4 vol %, and a combined propylene plus butylene yield of 35 vol %.

J. Williams: Light olefins can be increased by ZSM 5, high silica/alumina ratio zeolite, or low contact time. Other additives have yet to be as effective as the methods mentioned.

ZSM 5 has the advantage of increasing light olefins, even with a high activity catalyst, without increasing coke and dry gas. The olefin increase comes largely from mid range FCC naphtha of poorer octane.

High silica/alumina ratio zeolites increase light olefins through lower olefin saturation and cracking of heavy gasoline. There is also an increase in isomeric compounds in the gasoline due to high stability of secondary and tertiary carbenium ions.

Ultrashort contact time, and by that I mean 0.1 0.5 sec, provides the highest olefin ratios in light olefins and gasoline streams with olefinicities of 80%, or more in the butane/butene stream. This has not been a popular option due to the required increase in catalyst to oil ratio, catalyst activity, or reactor temperature to offset the loss of severity with ultrashort contact time.

T. Williams: We have used ZSM 5 seasonally in one of our FCCUs with the primary, interest of increasing refinery grade propylene for sales. One test run on this material showed a 1.4 liquid vol % increase in both C3 and C4 olefins with a corresponding decrease of 2.1 liquid vol % on our FCC gasoline. FCC gasoline octane did increase 1.1 numbers RON and 0.4 numbers MON.

Joseph B. McLean (Engelhard Corp.): Concerning light olefin yields, a lot of the choice oh which option to choose depends on which light olefins you want to make. For example, ZSM 5 is very much selective towards making propylene, relative to C4S, with 50 60 % of the yield shift typical), to propylene. So if you have a good outlet and a good value for propylene, that is a very favorable way to go.

Reactor temperature, again, is also somewhat nonselective because of the increase in dry gas make. On the other hand, if you are interested in making isobutylene, which is the primary, direction people are going these days for MTBE feed, we found that the combination of a selective low-hydrogen transfer catalyst at the minimum possible operating temperature that you can afford to work with within your constraints and objectives, is the best way to maximize isobutylene yield.

There were some details on these comparisons which were published in NPRA paper AM 92 45. Also, I do not think that ultrashort contact time by itself really does much of anything for olefin make. But what it does do is get you off of gas make limits, which allows you to increase operating severity through either temperature or catalyst, or some combination. The combination of those effects will allow you to make some pretty dramatic increases in potential light olefin yields.

G. Andrew Smith (Intercat Inc.): What we found is that, in most unit operations, you get equivalent amounts of propylene and butylene from the use of ZSM 5, and about 40% of that butylene yield is isobutylene.

We have also had people that have worried a little bit about the gasoline loss and have gone to a high rare-earth catalyst, using additional amounts of ZSM 5 to avoid the gasoline loss while increasing the octane and the olefins. Also the next generation of ZSM 5 additives are coming out now and they offer the possibility of increasing that isobutylene yield.

Joseph W. Wilson (Caltex Services (U.K.) Ltd.): I am going to have to take issue with the gentleman from Engelhard about the effect of low contact times. Hydrogen transfer reactions, which you are trying to suppress to achieve maximum olefin yields, are intrinsically slower than catalytic cracking reactions.

Anything that reduces the contact time will suppress the hydrogen transfer reactions that saturate your olefins. Consequently, short-contact time risers, good separation of the catalyst and the product vapors at the end of the riser, and, in fact, things like closed cyclones should all add to increase the olefin yields.

Joseph B. McLean (Engelhard Corp.): I agree that the reduction in hydrogen transfer is key to maximizing light olefin yield. However, the light C3 C5 olefins of interest here are not consumed by hydrogen transfer reactions; rather, their higher molecular weight precursors are saturated, resulting in gasoline range paraffins being formed and retarding further cracking to light olefins.

So simply reducing the contact time, by itself, will not increase these yields. The use of shorter contact times and such devices as closed cyclones typically work together with increases in catalyst to oil ratio and/or temperature to result in the observed olefin yield increases.

SULFUR CONTAMINATION

Has anyone successfully used in situ regeneration to regenerate steam methane reformer catalyst that has been contaminated with sulfur?

Laabs: We have successfully performed this type of regeneration. Sulfur is stripped from the catalyst by running a high steam to carbon ratio.

It is very important to bypass the low temperature shift reactor during this regeneration. The sulfur that is stripped will permanently poison the low temperature catalyst. Because of this bypass, we have found it necessary to run the reformed gas to the flare.

Rajguru: Our investigation indicates that in situ regeneration has been tried on occasion, with little or no success. We believe that a reliable and adequate sulfur removal system upstream of the reformer pays for itself by extending the catalyst and tube life. In situ regeneration also requires isolation of the reformer from the shift reactor to protect the shift catalyst.

Sloan: Sulfur contaminated steam methane reformer catalyst is not necessarily a serious problem. If the source of sulfur contamination is discontinued and the contamination of the reformer catalyst is not severe, the sulfur will be removed through normal operation.

It would be advisable to operate with a higher steam-to carbon ratio than normal, In some extreme cases, the sulfur has been burned off using air.

Occasional sulfur contamination of the reforming catalyst is easily reversed through normal operation combined with higher steam to carbon ratios to accelerate the sulfur removal rate. Many operators have used these techniques to remove sulfur in reformers. It is important to note, however, the sulfur will be carried downstream as it is removed from the reforming catalyst.

T. Williams: We did have a minor sulfur contamination problem, and we did a high-temperature steam strip. We were losing about 1 2% hydrogen purity, which we did recover.

Ken Chlapik (ICI Katalco): Many of our customers have successfully steamed both alkalized and non-alkalized reforming catalysts and have recovered activity by doing so. Much of the activity recovery will depend on the age of the catalyst and the amount of sulfur that has been incorporated on the catalyst. The stability of the catalyst support will also affect the success of the steaming. The stable support systems used in ICI Katalco reforming catalysts allow effective in situ regenerations to occur.

I reference the ICI Katalco paper "Safe Start up and Operation of Steam Reformers" by Dr. Ben J. Cromarty presented at the 1992 AIChE Safety in Ammonia Plants and related Facilities Symposium in San Antonio, Tex.

It is important to note that steaming will cause some sintering of the nickel sites on the catalyst, also affecting the activity of the catalyst. We normally recommend a steaming period of 6 12 hr at normal reformer operating conditions to achieve this kind of recovery.

Anders Nielsen (Haldor Topsoe A/S): In principle it should be possible to regenerate the catalyst by decreasing the sulfur in the feed back to nominal. However, while you can put sulfur on very fast, it only comes off at equilibrium level, and diffusion in the pore system further limits the rate.

When the catalyst is regenerated by steaming, sulfur is converted to H2S and SO2. The success of the regeneration in steam depends on the catalyst. If the catalyst is promoted with sodium or potassium, it is nearly unaffected, since alkali sulfates are formed.

Reference is made to an article by J. Rostrup Nielsen describing, the chemistry, and to a successful industrial regeneration in an article by Andriano and Sinaga. (References: J. Rosh Lip Nielsen, Catalysis 5, 1984, pp. 101 103. Ir. Andriano and O.G. Sinaga Ammonia Plant Saf. 32, 1992, pp. 232 235.)

CATALYST REPLACEMENT

Do you replace catalyst in diesel and gas oil hydrotreaters baed on pressure drop or activity (temperature)?

Bonelli: Our experience has been that our hydrotreating, catalyst replacement is most commonly done for recovery of activity, especially as long as the unit operating cycle has been reasonably, nominal. The proper design and operation and maintenance of feed filtration systems are critical, especially when you are charging feeds from storage or feeds from FCCU operations.

Brierley: The factors controlling catalyst change out timing are specific to each unit. In our case, the units were generally pressure drop limited for the first 12 years of operation.

Improvements to the fluid cokers and tank blanketing systems have reduced the level of contaminants in the hydrotreater feeds. As a result, our naphtha and heavy gas oil hydrotreaters are currently activity-limited.

Osborn: Depending on the design or operating limitation of the unit, the catalyst can be replaced, either due to the pressure drop limit of the reactor or end of run reactor temperature limit. It is important to select proper size (diameter) and shape of the catalyst to balance either limit.

Generally, hydrotreating reactors and their internals are operable with pressure-drop increases of 30 50 psi from start of run to end of-run. The end of run for low-sulfur diesel hydrotreating is in the 700 7200F. range.

Beyond this temperature, color degradation may occur. Once again, this is dependent on catalyst selection, feedstock, and design vaporization of the feedstock in the reactor bed.

Dilip David (Criterion Catalyst Co. LP): Many refiners use the method if top bed skimming when there is a pressure drop problem. This will allow the refiner to prolong the cycle until end of-run conditions are reached.

If there are pressure drop limitations and the source of the problem can be identified, then by correct size grading of the bed, the pressure drop problem can sometimes be alleviated.

Gordon Low (Unocal Corp.): If high pressure drop is limiting run length, you have to identify the cause of the high pressure drop. If there are filterable particulates in the feed, feed filtration will be effective.

You should also try, to minimize the storage of the feed, or store the feed only in floating roof or gas blanketed tanks.

If the pressure drop is caused by nonfilterable feed contaminants, then you may need to alloy the upstream units to remove soluble iron from the feed, or improve desalter and vacuum tower operation to improve the feed quality. In addition to these steps, which are designed to improve the feed quality, you can also extend run length by loading a graded bed in the lead reactor or installing a bypassing guard reactor.

REGENERATED HDS CAT.

Do you use regenerated hydrodesulfurization (HDS) catalyst? If so, in what hydrotreating service?

J. Williams: Regenerated HDS catalyst is used by most refiners in at least some services. Catalysts from distillate hydrotreaters and straight-run naphtha service are successfully regenerated in almost all cases.

Heavier feed units contaminated with metals or silicon at higher levels must be evaluated on a case by case basis. Typically, refiners would use regenerated HDS catalyst in units where the operation is not so severe or critical, such as straight run naphtha or kerosine hydrotreating.

Pedersen: For most of our hydrotreaters, we have found it economical to regenerate the catalyst off site two or three times before replacing the batch with new catalyst, even for our light cycle oil (LCO) hydrotreater, which processes a blend of light cycle oil and heavy coker gas oil. In our coker distillate hydrotreater, however, we have to replace the catalyst after each run, due to silica poisoning.

Bonelli: We cascade the hydrotreating catalyst from the guard bed of our hydrocracking reactor into naphtha hydrotreating service.

Brierley: We have used regenerated HDS/HDN (hydro-denitrogenation) catalyst in coker naphtha service, coker gas oil service and for combined hydroprocessor and virgin gas oil. We have traditionally used second and third cycle catalysts in those units that were pressure drop limited.

Now that some of these units are activity limited, the economics of using higher-activity, catalyst, including fresh catalyst, are being evaluated.

O'Brien: We have used regenerated HDS catalysts in naphtha, kerosine light cycle oil, distillate, and residual service. In our next residual hydrotreater catalyst load, we are going to use some regenerated distillate HDS catalyst in the last reactor.

Osborn: I generally agree with the cascading concept but I would caution that if you are going to use it in a naphtha service, you have an expensive Platformer catalyst following that, so you need to make sure you have good catalyst and good activity.

Selecting regenerated catalyst is, certainly, dependent on it being free of contaminants and you need good recovery of activity (80 85 % minimum) of surface area and pore volume vs. fresh catalyst. We have purchased regenerated catalyst for use in kerosine, coker cycle oil, and coker naphtha service, with good success.

SOLID ACID CATALYST

What is the status of solid catalyst alkylation research? What is the possibility of the development of a commercial process and what is the expected time frame? Is the reactor a fixed bed or some type of continuous regeneration system?

Juno: Neste Oy, Catalytica Inc. of Mountainview, Calif., and Conoco Inc. started up a 7 b/d pilot plant at Neste's Finland technology center to demonstrate solid catalyst gasoline alkylation technology.

Their proprietary process is being developed to replace hydrofluoric and sulfuric acid alkylation units used to produce high octane gasoline. Licensing is to be available this in 1994.

Sloan: Haldor Topsoe A/S and M.W. Kellogg Co. have been collaborating to develop an alkylation process which is a substitute for both the current HF and sulfuric acid processes.

The "super" acid catalyst is supported on a solid medium in a unique manner. The reactor system can be operated continuously without interruption for catalyst regeneration.

Acid inventory is very low and acid management is greatly enhanced, since the catalyst can be recovered and recycled on the refinery site. The catalyst also has a low vapor pressure and there is no tendency to form an aerosol.

The technology has been successfully tested in a pilot plant of 0.5 b/d for over 2 years, including a continuous run of 6 months. The technology is being commercialized and discussions with potential operating partners are in progress.

Brian Johnson (UOP): As the industry is aware, UOP has been working for a number of years to develop a solid catalyst alkylation process. At this point, UOP has developed a truly solid phase regenerable catalyst and is moving into long term stability tests in our pilot plant facilities near Chicago.

An initial engineering package is being developed that will be competitive with the capital and operating costs of current technologies.

UOP will present the technology to potential licensees in 1994 and expects to have a complete commercial design package by 1995. UOP will continue to support the current HF technology with ongoing developments, such as the Texaco/UOP HF additive technology, as well as minimum acid inventory designs, redundant acid storage designs, and rapid acid dump systems.

Lawrence A. Smith (Chemical Research & Licensing Co.): Chemical Research & Licensing has announced a solid acid alkylation process based on a "salt" of antimony pentafluoride. A 10 b/d test unit will be in operation at Chevron, Port Arthur, Tex., before the end of 1993. A 1 year test program is planned on a variety of commercial feedstocks.

HYDROGEN YIELD

What gains in reformate and hydrogen yields do refiners typically achieve on semiregenerative reformers using bimetallic catalyst when reactor pressure is reduced by 100 psi?

Juno: In 1992, we revamped our semiregenerative reformer, lowering the separator operating pressure from 300 to 170 psig. Based on prerevamp and postaudit detailed test runs, we achieved an increase in C5+ yields of 2.7 liquid vol %, and an increase in hydrogen production of 210 scf/bbl.

Measured research and motor octanes were within 0.2 of an octane number and the (R + M)/2 was within 0.1 of an octane number, between the two test runs.

Projected yields from UOP for the R62 catalyst for separator pressure reduction from 300 to 200 psig were a 2.1 liquid vol % C5+ yield increase, and 164 scf/bbl hydrogen production increase. Projections of catalyst cycle length between regenerations were exceeded.

O'Brien: Although somewhat the reverse of the question, in the case of one of our motor fuel reformers, we increased the reactor pressure by 140 psig. During the test period the unit was processing feed with N + A (naphthenes and aromatics) of 60-66 vol % and operating with seventies ranging from 91 to 99 C5+ octane.

The following was observed: hydrogen yields decreased 0.19 wt % for the low severity case and 0.3 wt % for the high severity case; the C5+ liquid yield decreased 1.0 vol % for the low-severity case and 0.5 wt % for the high severity case; and the catalyst stability was increased approximately 20%.

Bonelli: At one of our refineries, when we lowered the pressure on a semiregenerative unit by 80 psig, we realized a gain of 2 3 vol % in reformate yield, and an increase in hydrogen production of 50 100 scf/bbl.

Paules: For a typical semiregenerative reformer operating at an average reactor pressure of 300 psig and 95 RON severity, a drop of 100 psi average reactor pressure will yield an increase of approximately 1.0 vol % reformate and 100 scf/bbl of hydrogen on a rich 65N + 2A feed, and an increase of approximately 2.2 vol % reformate and 150 scf/bbl of hydrogen on a lean 40N + 2A feed.

Pedersen: From our pilot plant, we have found that reducing the reactor pressure from 375 to 275 psig increases C5+ yield by as much as 5.0 wt %. The corresponding hydrogen yield will increase by approximately 35%.

Feed in our test were naphtha from a 40/60 mix of Statfjord and Gullfaks crudes. Tests were carried out over a severity range from 96 to 102 RON clear, with a platinum/rhenium catalyst.

M. Edward Morrison (IFP Enterprises Texas Inc.): With a semiregenerative reformer using bimetallic catalyst and operating at 250 300 psig (mean reactor pressure), the estimated yield improvement at constant octane, after reducing pressure by 100 psi, is estimated to be 2.5 wt % for C5 liquid yields, and 250 scf/bbl for hydrogen yields. This obviously reduces the cycle length of the semiregenerative catalyst.

Lee E. Turpin (Profimatics Inc.): The changes in reformate and hydrogen yields as pressure is changed vary significantly with feed composition. The changes in C5+ and hydrogen yields decrease as the quality (N + A) of the feed increases.

The accompanying table was generated using Profimatics kinetic reformer model. C5+ octane was held constant as well as feed boiling range.

One of the reasons I have seen people raise pressure is because the are meeting heater constraints, and if the raise the pressure, the can have a bigger heat sink with the hydrogen and reduce the heater constraints on downstream heaters.

SULFATE FORMATION

What is the effect of sulfate formation during the regeneration of catalytic reformer catalyst? Can the catalyst be restored to fresh catalyst activity and stability conditions? What are the methods used to correct the sulfate problem?

Osborn: The existence of sulfate, SO4, on catalyst during a regeneration has two serious effects. First, the sulfate promotes platinum mobility. In turn, this leads to platinum agglomeration and loss of catalyst stability. Second, the sulfate hinders the ability of the catalyst to pick up chloride. This in turn leads to loss of catalyst activity.

Both effects also result in a loss of yields. As long as the sulfate is present, the platinum crystals cannot be properly dispersed.

The catalyst performance can be restored to fresh catalyst activity and stability through an effective sulfate-removal procedure. This involves removal of the sulfate in a hydrogen atmosphere with a large presence of chloride. Chloride competes with the sulfate for sites on the alumina, thus driving the sulfate off the catalyst. The sulfate is removed as H2S.

Laabs: We agree. In addition, we recommend sulfate removal if there is more than 0.05 0.1 wt % sulfate on the catalyst.

Rajguru: I agree with the previous comments on the first part. The sulfate on the surface of a reforming catalyst prohibits, or at least decreases, the rate of noble metal redispersion. Yes, the catalyst can be restored to close to fresh activity when carrying out a sulfate sweep during the hydrogen treatment between the carbon burn and the oxychlorination step.

Sulfate can be effectively stripped from the catalyst by circulating relatively high purity hydrogen at elevated temperature. Chloride addition will improve the rate of sulfate removal.

J. Williams: Sulfate formation will make the redispersion of the platinum more difficult at the time of the oxychlorination and can cause sintering at the time of reduction because of the decomposition of the sulfates into water and H2S. The result is a very low activity after regeneration. We were able to restore catalyst performance by double oxychlorination with intermediate H2S stripping. A caustic wash circulation is necessary in the H2S stripping stage to absorb the formed H2S from the recycle gas and prevent recontamination of the system upstream of the catalyst.

T. Williams: We saw evidence of sulfate formation during a regeneration that followed a sulfur excursion. We did extensive hot hydrogen stripping according to UOP procedures. Typical start of run activity was achieved.

Thomas W. Kelly (ARCO Products Co.): I agree that full catalyst activity can be restored if sulfur is completely removed before the chloro-oxidation step. High sulfur on catalyst during this step will cause platinum crystallite growth to such an extent as to make good metal redistribution not possible.

A study in our process lab confirmed that nothing could restore catalyst to new activity after it had seen a reactivation while contaminated with 0.5 wt % sulfur. The catalyst had been through a 12 hr preregeneration hydrogen strip to remove sulfur, which was obviously not very effective.

Himmat Singh (Indian Institute of Petroleum): Due to some plant upsets, we did experience sulfate formation which ranged between 0.32 and 1.65 wt %. The catalyst in this case was regenerated in two steps, following the procedures outlined by the panel. The activity of the regenerated catalyst was close to its original value.

Terry Tucker (Mapco Petroleum Inc.): We have had success with the sulfate stripping procedure, as recommended by UOP for R62 catalyst. As you are doing a regeneration burn, the SO2 appears to accumulate in the reactor bed unit near the end of the burn. Therefore, you need to test for SO2 via Draeger, just as oxygen breakthrough occurs at the end of the burn.

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