NPRA Q&A--1 PRESULFIDING, FCC OLEFINS FOCUS OF CATALYST SESSIONS

March 16, 1992
Increasing demand on refinery operations spurred much discussion of refining catalysts and additives at the National Petroleum Refiners Association annual question and answer session on refining and petrochemical technology, held Oct. 16-18, 1991, in Denver. Herbert Bruch, NPRA technical director, moderated the renowned meeting, where representatives of the international refining and petrochemical industries come together from around the world to discuss pertinent and diverse topics.

Increasing demand on refinery operations spurred much discussion of refining catalysts and additives at the National Petroleum Refiners Association annual question and answer session on refining and petrochemical technology, held Oct. 16-18, 1991, in Denver.

Herbert Bruch, NPRA technical director, moderated the renowned meeting, where representatives of the international refining and petrochemical industries come together from around the world to discuss pertinent and diverse topics.

Presubmitted questions are first answered by a select panel of refiners, petrochemical producers, and industry experts (see box). Then the audience is invited to respond with comments or additional questions.

This first abstract from the recently released transcripts of the Q&A session addresses the role of catalysts in fluid catalytic cracking (FCC), hydrotreating, and hydrocracking processes.

FCC OLEFINS

Some of our laboratory studies have indicated that a ZSM-5 additive does increase C3/C4 olefins production, but it appears to be highly selective toward the propylenes, whereas our processing constraints call for more butylene production. Has this selectivity been observed in commercial operation with ZSM-5?

McClung: Engelhard is monitoring several units using our ZSM-based additives Z100 and Z1000. These units vary the additive addition rate to obtain the desired octane level for the FCC gasoline, as well as to get increased butylene yield for MTBE production or inCreased total butylenes for the alkylation unit.

Upon addition of ZSM-5 additive, we observed that 50-60 vol % of the incremental increase in light olefins, that is the C3s and C4S, was propylene. The ratio does change as a function of the feed quality and reactor temperature.

Pazmanyi: We had a 1-month-long additive trial in 1990. In a 2-week time, we reached 4% additive level on the catalyst. The effect was quite obvious, with a stepwise rise of the effect of the additive. After stopping the additive injection, the effect was fading down in a 10-day period of time.

The measured yield changes in percent on feed were, in the case of C3 unsaturates, +0.87%; and for C4 saturates, +1.35%. The research octane number (RON) increase was 0.7; the motor octane number (MON) increase was 0.9. The C4 olefins breakdown can be interesting from the point of view of methyl tertiary butyl ether (MTBE) production. Our analysis shows that the rate of change for iso and normal butylenes is the same.

Williams: We have had several experiences with ZSM-5 on both of our FCC units. Our commercial data show a high degree of selectivity towards C3 olefins over butylene production. Incremental propylene yield exceeded butylene yield by about three to one on a volumetric basis.

Daniels: We also have seen higher yields of propylene than we thought we would have at conventional concentrations of ZSM-5 in our unit.

We were able to significantly increase the ratio of butylene to propylene yields by significantly increasing the concentration of ZSM-3.

Fischer: We have experience with ZSM-5 in two of our refineries and have generally seen a higher selectivity to propylene similar to the other panelists. We have just completed a trial at one refinery. The vendor predicted an increase of 0.7 vol % in propylenes and 0.6% in butylenes. During the trial we actually saw a 2% propylene increase and a 1.3% increase in butylenes. Gasoline yield declined by 3 vol % with a 0.5 RON increase. Feedstock changes may have played a role in these shifts, and this is still being analyzed.

Grant: One of our plants which uses ZSM-5 has observed a 0.2 wt % increase in propylenes for every 0.1 wt % increase in butylenes.

Daniel J. Neuman (Katalistiks International Inc.): Mr. McClung mentioned that certain changes in conditions, and I believe he said feedstock, will change the ratio of C4 olefins to C3 olefins. Can you expand on that please?

Anthony Witoshkin (Engelhard Corp.): In commercial operation, we look at two other parameters when comparing incremental C3 C4 olefins ratio changes: reactor temperature and feed quality. The latter has a greater impact since it affects yields and the ratio of linear olefins, paraffins, and alkyl aromatics in the fluid catalytic cracking unit (FCCU) products. Reactor temperature also affects these to a lesser degree. Cracking of these hydrocarbons via ZSM-5 will impact C3-C4 olefins ratio. The ratio will depend on the molecular weight of linear hydrocarbons and the side chain length of alkylaromatics.

Hartley Owen (Mobil Oil Corp.): We have observed the same thing commercially that the panel has observed. The only thing I would like to say is that we believe the selectivity can be modified.

David Roberts (Amoco Oil Co.): In your commercial tests, at what base loading amounts did you operate the unit? Also, if you ran several tests at different base loadings, what effects were observed?

Daniels: We started our trial out at 1.5% of ZSM-5, and that is where we saw our higher-than-expected propylene yields. Through time, we Eventually built that concentration up to 6%, and saw a real large shift back towards the butylene and away from the propylene.

Steven Nelson (Kerr-McGee Refining Corp.): Do you have any comments on the amylene increase relative to the C3/C4 olefins?

Pazmanyi: I have one data point concerning the C5 olefins. There is an increase of 0.05%.

Daniels: Our commercial results have shown that the percent increase in amylene production is very close to the percent increase in the butylenes.

William R. Reid (W.R. Grace & Co.): "le have also seen a selectivity preference for the propylenes. Our commercial experience would indicate that we expect to see about 55 vol % C3 olefins in the incremental LPG olefin production. This is also a function of the amount of ZSM-5 that is used in the unit. As the level of ZSM-5 goes up, you will see an increasing ratio of C3 olefins. That reflects the selectivity of the ZSM-3 zeolite.

ALKYLATION-FEED

What FCC catalysts maximize alkylation feed and how much alkylate can be made by this route?

McClung: The catalytic properties that increase gasoline octane will also maximize alkylation unit feed; the C3 through C5 olefins. This catalyst needs to be a USY zeolite with an active matrix. Zero rare earth USY will minimize hydrogen transfer reactions, thereby maximizing olefin yields. An active matrix will complement the zeolitic cracking by providing additional low hydrogen type of cracking on its surface. Inclusion of ZSM-5 type additive will further boost the C3 through C5 olefins production.

The chances in the catalyst types from high rare earth, low inactive matrix surface area to zero rare earth high and active matrix surface area can increase propylene and butylene Production to alkylation units by 30-50%.

Daniels: We have taken a look at this with two different types of feedstocks: high-nitrogen and low-nitrogen feeds. With a low-nitrogen feed, a high activity, zero rare earth catalyst with ZSM-5 makes the most alkylate. With a 10% inventory of ZSM-5, our laboratories have shown a 50% increase in alkylate feed as C4 olefin. On higher nitrogen feedstocks, we have found that zero rare earth catalysts cannot hold conversion, so that ZSM-5 should be used in combination with a high rare earth catalyst to maximize alkylate feed. Using 10% ZSM-5 and a high rare earth catalyst on a high nitrogen feedstock can increase alkylate feed by about 25%.

ISOBUTYLENE PRODUCTION

Have any refiners changed or modified the catalyst they are using in their FCCUs to produce more isobutylene for their MTBE unit? With what success?

Daniels: Catalytically, there are two ways to increase isobutylene production. One would be to increase your total butylene production at a constant isobutylene-to-butylene ratio. The other way is to increase your ratio of isobutylene to total butylene. Total butylene production can be increased catalytically by decreasing unit cell size, increasing catalyst activity, increasing matrix surface area, or adding a ZSM-5 additive. Commercially, we have increased our butylene production over 10% using ZSM-5.

To date, the only catalyst systems that we have seen in the laboratory that significantly increase the isobutylene-to-total butylene ratio are systems with ZSM-5. We have increased the isobutylene-to-butylene ratio from about 30% to 40% when we use very high levels of ZSM-5 additives. In our unit we have used as much as 6% ZSM-5, and we have not been able to confirm an increase in isobutylene-to-butylene ratio.

McClung: Engelhard is aware of a unit that uses varying amounts of ZSM-5, in Combination with operating conditions, to adjust isobutylene yield to suit the MTBE unit requirements. They have been practicing this for some time and are quite satisfied with the results and the operational flexibility then, get.

Ales Soudek (Pace Consultants Inc.): I have a couple of points to make here. First, there is a thermodynamic ratio of isobutylene to butylene that can be attained, and that is about 42-44%. Secondly, there was a paper presented in the 1991 NPRA annual meeting which talked about a new catalyst additive which had the same effect as ZSM-5 with the exception of producing predominantly more butylene and amylene yield, and less propylene yield. I wonder if anyone has used that in commercial operation or whether there was a pilot plant test?

Kent Davis (W. R. Grace & Co.): It is catalytically possible to affect the yield of isobutylene from the FCCU. However, it is also important to consider the impact on other FCC yields as a result of doing so. We have some experience with our XP technology catalysts in which we have been successful in increasing the isobutylene yield and isobutylene selectivity from the FCCU with no loss in gasoline volume.

CATALYST PRESULFIDING

What has been the industry's experience with using ex situ presulfided catalyst? Any problems with equipment plugging downstream of the reactor? Any material handling problems? How much did it shorten start-up time? What steps make up the start-up procedure with this catalyst?

Moyse: Our experience with the use of ex situ presulfided catalyst has been good, both in naphtha and gas oil service. We have so far seen no evidence of plugging or handling problems. We have experienced approximately 50% reduction in time required for start-up, i.e., 1 day as opposed to 2 days for the traditional presulfiding using dimethyl disulfide (DMDS) or native sulfur.

The start-up procedure is very simple. You heat up from ambient to 625-660 F. at a rate of 100 F./hr or as fast as your equipment allows. An exotherm of about 60 F. will be observed around 320 F., and heating up may at this stage be stopped in order to avoid a too-high reactor outlet temperature. Recycle of H2S-containing gas and liquid is strongly recommended to avoid loss of sulfur during the start-up.

Pazmanyi: We have used in recent years mainly presulfided catalyst, for naphtha hydrotreaters and gas oil desulfurizers, from Akzo. Downstream of the reactor we did not experience any equipment plugging. We do not have handling problems; the works inside the reactor (like catalyst loading and installation of ceramic balls, baskets, etc.) are done using breathing apparatus and safety clothes.

Our experience shows that the start-up time can be shortened by 1 day on average. Another advantage of this presulfided catalyst is that we do not have to handle DMDS or other toxic materials.

I think the start-up steps vary depending on the catalyst and the philosophy of catalyst manufacturers. Usually exotherms can be expected at certain temperature levels. Our experience is positive in respect to presulfided catalysts. We see better activity and better overall performance.

Pritzel: The experience I am familiar with has been good. I am not aware of any problems with downstream plugging. The material can be handled with or without an inert atmosphere.

By eliminating the presulfiding step it saved approximately 1-1/2 days.

Knepper: We are kind of the lone dissenter here. We had a load of ex situ regenerated catalyst presulfided by the regenerator. A large reactor outlet temperature was observed upon start-up and catalyst activity was very poor thereafter. We felt that the 10-hr reduction of start-up time was not worth the catalyst changeout that was required a month later.

Frederickson: Chevron's only experience with presulfided catalyst was related at last year's NPRA Q&A session. Our comments, along with those of several other refiners that have tried presulfided catalyst, are on pages 98-99 of the 1990 NPRA Q&A session transcript. Most of the industry's experience has been better than ours.

Fischer: Two refineries in the Sun system have had experience with ex situ presulfided catalyst. In April of 1989 our Toledo refinery sent out 24,000 lb of naphtha hydrodesulfurization (HDS) catalyst for regeneration and presulfiding. In July of the same year we loaded a new catalyst charge into a kerosine HDS unit in our Sarnia refinery. This catalyst had also been presulfided.

Start-up of the presulfided catalyst went well in both cases and both units are performing well. In 1988 we experienced a problem, self-inflicted, with ex situ presulfided catalyst in another one of our refineries. The catalyst was exposed to air during storage in super sacks, resulting in some heat generation and gas evolution.

The response from the manufacturer was as follows:

  • Updated labeling and MSDS to include handling precautions similar to those for spent catalyst.

  • Storage in containers other than sealed steel drums with plastic liners is not recommended for more than 16 hr prior to loading.

  • Exposure to heat is definitely to be avoided.

Our general feeling, however, is that shortened start-up time is offset by the additional cost and special handling procedures required for safe operation. Ex situ sulfiding is advantageous in specialized cases such as with reactors with poor liquid distribution that could lead to incomplete in situ sulfiding.

Golden: One of our refineries uses presulfided catalyst in the naphtha hydrotreater, saving about 1 day in the start-up procedure. The fresh catalyst is heated to 250-300 F. at normal pressure with once-through hydrogen. Activation of the catalyst is indicated by an exotherm. Once the exotherm subsides, the catalyst is heated to 400 F., and feed is introduced. No particular problems have been encountered with this procedure.

Ted W. Mole (Conoco Inc.): We have also had very good experience with presulfided catalyst. I want to note that usually plugging will only occur when you have a presulfided catalyst made by merely placing the elemental sulfur on the surface of the catalyst. There are other catalyst manufacturers who will actually bond the sulfur to that catalyst.

After three successful presulfided catalyst start-ups, in March 1990 the Denver refinery was the third refinery to experience feed effluent exchanger plugging during start-up with presulfided Criterion catalysts Similar to other plants' we plugged at reactor bed temperatures of approximately 250 F., which is near the melting point of elemental sulfur. Apparently sulfur was slightly soluble in the distillate and washed off into the feed effluent exchangers where it deposited and solidified due to the slightly cooler feed effluent exchanger temperatures.

Unlike our refinery in Billings, Mont., we were able to work our way out of the problem within about 45 min by maintaining heat and pressures We believe that our 100% recycle feed and hydrogen lineup helps keep these exchangers warmer than with a once-through lineup. In order to prevent problems we recommend four procedure changes:

  1. Maintain a lower feed rate sufficient to remove heat but minimize the amount of sulfur washed off.

  2. Maintain a higher hydrogen flow rate to help push the reactor without solubility concerns. We use approximately 3,000-5,000 b/d distillate feed and 5,000,000 scfd hydrogen, which is 25% of our normal feed rate and 40% of our normal hydrogen rate.

  3. We also maintain that you should have an aggressive heat-up rate of approximately 100 F./hr through 300 F.

  4. It is also very important to maintain a stable flow feed rate through the 200-350 F. temperature range. We also use 100% recycle and we would recommend that as well.

I will also relay that we had an experience with poor catalyst activity on our March 1991 start-up. While we cured the feed furnace refractory we circulated nitrogen throughout the furnace tubes and the freshly loaded reactor. The nitrogen temperatures reached approximately 400 F. We suspect that the sulfur flowed off the catalyst surface during the week-long refractory cure-out.

I will also relate that we started up our unit with a new CRI presulfided Criterion catalyst in July. The new presulfiding technique used by CRI provides a much improved product and an activity that is approximately 50% higher than previous products and a 30% longer estimated run length than we had with the previous lot of presulfided Criterion catalyst. The only difference during start-up, beyond eliminating sulfur wash-off concerns, is that the exotherm occurs at 500 F. rather than 350-400 F.

Mark J. Sally (Conoco Inc.): We have talked about the presulfided catalyst; I am wondering why the catalyst manufacturer; suggested maintaining a reactor delta T of less than 25 F. during presulfiding when using a sulfiding agent, and there seems to be no delta T limit when using a presulfided catalysts.

McClung: I believe I can answer that question. Does your supplier indicate that it is true for both cobalt-molybdenum and nickel-molybdenum? Because with nickel-molybdenum you would be concerned about nickel reduction before you got it sulfided. But after it is sulfided there is not as much concern. That is the only reason I can think of.

Mehmet Y. Asim (Akzo Chemicals Inc.): I just have a few additional comments and a point to clarify what has already been mentioned. The supplier of the ex situ presulfided catalyst Mr. Fischer referred to is Akzo, and until recently we had been specifying shipment and storage of our EasyActive catalysts only in steel drums with sealed plastic liners. However, since the summer of 1991, Akzo has been delivering the presulfided catalyst in flow bins (87 cu ft capacity each), supplied by Federal Container Corp. As such, shipment and storage in drums only is no longer a limitation.

As for Mr. Sally's question, in situ sulfiding spike agents are done at temperatures of 450-530 F. At these temperatures, a conservative delta T of 25 F. is strongly recommended so as not to exceed an outlet temperature of 575 F. Otherwise, in the presence of hydrogen, the oxidic metals of the catalyst would be reduced to their lower valence states rendering them much more difficult to be sulfided.

This is not a problem at all with our ex situ presulfided catalysts since the sulfur is fixed on the catalyst as oxy-sulfide species, which only require activation at 300 F. to be converted to the final metal sulfides. Due to this low activation temperature, much higher reactor delta Ts can be tolerated with Easyactive catalysts.

J.B. Roddey (Roddey Engineering Services): No one has mentioned this but, although you save time with presulfided catalysts, you also limit yourself on flexibility of start-up procedure. line of the key issues in having good catalyst activity is having the catalyst at a very dry state before the sulfiding is started. With the presulfided catalyst, you do not have much flexibility on the mechanisms by which you dry the catalyst prior to starting the sulfiding. I can assure you it will make a major difference in activity after sulfiding whether the catalyst is wet or dry.

Art Suchanek (Criterion Catalyst Co. L.P.): This is a very simple answer to a 25 F. delta T. When you are bringing in sulfur to a catalyst that is in an oxidized state, you get a flame front pickup of the H2S or of the sulfur, whereas the presulfided catalyst has sulfur all through it. You will get the chance of the sulfiding being taking place in that way.

James J. Barry (CRI International Inc.): I do not want to belabor the point but I would just add a few comments. The Chevron experience was on CRI catalyst and it was a bad experience. We went back to the laboratory and have subsequently developed an improved modified material which has eliminated the sulfur transport problem. We have had a couple of commercial trials, one of which you heard about from Mr. Mole of Conoco. We expect to offer this in the future. Right now you can get it on regenerated catalyst with CRI or on fresh catalyst with Criterion.

AROMATICS SATURATION

What catalyst developments have been made to improve aromatic saturation in hydrotreaters?

Hamilton: The application of combinations of highly active non-noble and noble metal catalyst at low operating pressures in the range of 750-950 psi has made it possible to achieve diesel product with aromatics content of 10% or less by FIA. This synergetic catalysis approach, as developed by Art Suchanek of Criterion Catalyst, presents refiners with the opportunity to utilize the proper catalyst or combination of catalysts for meeting the October 1993, diesel desulfurization regulations, while providing the refining industry the ability to produce low-aromatics diesel in the future.

Hunter: Unocal has supplied a high-activity aromatic saturation catalyst called AS-100 for many years. This catalyst was originally designed to upgrade turbine stocks at low pressure and it has been used commercially in many units. Over the last few years, Unocal has done extensive process research to expand the application of this catalyst to diesel boiling range stocks including cycle oils at pressure levels below 1,000 psi. We think that a good option is to treat some of your streams to very low levels of aromatics and blend with the untreated streams. This can effectively minimize the size of the necessary equipment.

McClung: Base metals, meaning nickel-based catalyst, nickel, nickel-molybdenum, or tungsten ind noble metal-based hydrogenation catalyst, have been known to the industry, both oil and chemical, for decades. Only in recent years has the deeper saturation, that the lubricants industry and the fine chemical industry have used for years, been required for deeper saturation of aromatics.

Therefore, the most significant improvements for deep saturation had been in using these types of catalysts in new process applications, most of which have been eloquently described by various vendor spokesmen today. It is my own personal belief, the most significant innovations in terms of development of catalyst have been in the nickel-tungsten area.

Moyse: We agree with Mr. McClung. There are several options open for aromatic saturation based both on new catalyst developments and also conventional types which we are all familiar with.

For moderate reduction of aromatics, say from 35% in the feed down to 25% in the product, a two-stage process using a highly active nickel-molybdenum catalyst in the first stage followed by a nickel-tungsten catalyst in the second is probably the most feasible.

For deeper aromatic saturation where more than 50% aromatics removal is required, the best option will be a two-stage process with a sulfur-tolerant catalyst in the second stage. These catalysts, which represent a new development in aromatic saturation, can tolerate up to 500 ppm of the sulfur in the feed going to the second stage.

Pazmanyi: At the 1990 NPRA annual meeting and the 1991 Akzo Symposium, papers on "Hydrotreating for Ultra Low Aromatics in Distillates" and "Aromatics Hydrogenation of Diesel Feedstock" were presented, which gave some interesting information about this topic.

Also, I would like to say that we are developing our own technology where, using Hungarian low-sulfur gas oil, we can produce aromatic-free or low aromatic content product which could be used in heavily loaded, densely populated areas, for example, Budapest. The process comprises two steps: one is the desulfurization using CoMo catalyst and the second step is the aromatics removal using platinum-containing catalyst at the lower temperature.

We can achieve, according to our laboratory, results, less than 1% aromatics content. If we get help in financing from our Minister of Environment Protection and we can realize an adequate price for this product, we would like to go ahead with this project.

Richard Foley (Zeolyst Enterprises): Another significant advance has been made with the Zeolyst 714 catalyst series. One of the problems at 500-600 psi of hydrogen partial pressure is that at 600-625 F. you will reach aromatics equilibrium. Of course you cannot reduce aromatics below equilibrium.

The thing to do is defeat equilibrium by ring opening very selectively the di and trinaphthenes. Z-714C will do that without ring opening the mononaphthenes. So you can, in essence, beat equilibrium by eliminating the reaction product. Very little gas and light products are formulated using this sulfur-tolerant noble metal on zeolite aromatics saturation catalyst (Z-714C).

HYDROCRACKING

Could you discuss the merits and drawbacks of using NiMo vs. NiW vs. Pd-based USY catalysts for fuels hydrocracking?

Moyse: The merits and drawbacks of these catalysts depend mainly on two things. One is the processing objective, i.e., whether the refiner wants to maximize the yield of naphtha or maximize the yield of middle distillate; and two, the contaminants, i.e., sulfur and nitrogen in the feedstock. All of the catalyst mentioned are bifunctional. That is, they have an acid function and a hydrogenation function.

The acid function is associated with the zeolite support and increases with increasing zeolite content. The hydrogenation function is associated with the metal or metal sulfides used.

Hydrogenation activity is of three types; palladium is best, nickel-tungsten is second, and nickel-molybdenum is third. The hvdrocracking activity increases with an increase of both functions of the catalyst. As hvdrocracking activity of the catalyst increases, so too does the selectivity to light distillates. Thus, palladium-based USY high zeolite catalysts are preferred for maximizing gasoline yield. Nickel-molybdenum or nickel-tungsten catalyst on low zeolite-containing supports are better for maximizing middle distillates.

Nickel-molybdenum and nickel-tungsten exhibit higher sulfur tolerance and are better suited for the higher sulfur-containing feedstocks. For a given conversion level, palladium-based USY catalyst can be operated at a lower temperature than nickel-molybdenum or nickel-tungsten catalyst. Finally, palladium catalysts also give a lower gas make.

Richard Foley (Zeolyst Enterprises): If you look at the difference between the nickel-molybdenum, nickel-tungsten, and the palladium-based zeolite catalyst, it is a lot easier to start up, a lot easier to regenerate, and it is a lot less expensive to buy the nickel-tungsten or nickel-molybdenum variety. The difference in activity between the noble metal and non-noble metal is a function of how you are operating it.

At very high H2S partial pressures you can totally eliminate the activity differences between the noble metal and the non-noble metal catalysts. By eliminating the temperature you are also eliminating the selectivity differential. At low levels of H2S, you will saturate a lot of aromatic products and use up a lot of valuable bad temperature differential.

Lastly, on the difference between nickel-molybdenum and nickel-tungsten, if you optimize the catalyst manufacturing procedures, there is no significant difference in activity or selectivity. The type of zeolite used is much more important. Product selectivity is determined by the catalyst support, not by the metal, e.g., zeolite type, ASA, alumina.

CATALYST REGENERATION

Does any refiner have experience with switching to a lighter or more-easily hydrotreated feed to a hydrocracker prior to a turnaround and off site catalyst regeneration? What feed and/or operational changes were made? What was the impact on the commercial regenerator's throughput rate and cost? How was the quality of the regenerated catalyst affected?

Williams: Prior to unit shutdown involving removal of catalyst from the reactor, we switch to virgin diesel oil as feed for the gas oil HDS unit. Diesel oil is brought into the unit and circulated throughout the reactor loop and fractionation section. The hydrogen circulation rate is maintained and reactor inlet temperature is reduced to approximately 500 F.

The benefits include: a more uniform catalyst bed cool-down process, vastly easier removal of the hydrotreating catalyst, which translates to a time savings; time saved during the regeneration process due to cleaner catalyst; and field instrumentation such as float chambers and sight glasses are much cleaner and easier to work with.

Moyse: I only have part of an answer. We agree that during shutdown of a unit treating heavy oil it is advantageous to introduce a lighter oil to wash out the heavy oil and any other contaminants that you can from the catalyst system. We have no idea what impact this has on regenerability.

Grant: At one location, our normal procedure for hydrocracker shutdown is to switch from gas oil to a 100% straight-run diesel feed for a 24-hr period before shutdown. To accommodate the change in feedstock, the charge rate is reduced by 4,000-6,000 b/d, and reactor inlet temperature is reduced by 20 F. By incorporating the change in feedstock to the shutdown procedure before the hydrogen strip step, lower explosion limit problems during unloading are alleviated.

In another facility, the feed is switched from a heavy, gas oil to a light gas oil. The reactor inlet temperature is reduced from about 800 F., to 650 F. during the changeover in feedstock. The light gas oil is run through the reactors for 20 hr. After 20 hr, the feed is pulled out of the unit and hydrogen-stripped step at 750 F. is started. The hydrogen-stripped step is done for 12 hr.

We just have qualitative data on the commercial regenerators throughput. We understand the regeneration has proceeded much more smoothly and the time for regeneration has been significantly reduced.

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