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OHMES (KBC Advanced Technologies, Inc.)

Again, a lot of these comments are in the Answer Book, so I will skip through part of my remarks. People often forget about the refinery slop processing system. Typically, refiners understand where their cracked stock goes; but if it is typically or abnormally routed to the crude tanks or units, one will have cracked stock throughout the refinery, which can cause problems.

Some of the mitigation options can be summarized in four areas. The first is feed logistics. We would say that if you are going to handle this, you need to do as much direct feeding as possible. Again, it is a refinery; so, it is really not practical to be 100%. If you are going to use storage, all of the tank streams will need to be under a nitrogen blanket or a floating roof tank utilized, even if these are virgin streams that will mix with cracked stock. You will obviously need to check your tanks during regular maintenance and cleaning because, again, the primary causes and ingredients of this foulant formation are the olefins, diolefins, the iron scale around it that acts like a catalyst, and the oxygen.

Once that is under control, another option would be chemical treatment. Some people do put in oxygen scavengers; they will work. Others put in different inhibitors to stop the polymerization from forming, particularly on the initial exchangers. Again, there are mixed results in the plants. Some work; others do not. It is probably worth a trial to see if it fixes your situation.

An important area of consideration is the unit configuration. The traditional example is determining how to handle coker naphtha. Some people will have the low temperature diolefin saturation reactor upfront, which can very effective. For those who do not, sometimes putting this coker naphtha in other spots (the response given earlier about putting it in the ULSD or the gas oil unit) is an option. In Asia, many refiners actually put in oxygen strippers to help manage this issue. They will bring in fuel gas, hydrogen, nitrogen, or other media to strip out the tankage feed to remove the oxygen.

The final recommendation concerns the catalyst; and, again, I will defer to the catalyst vendors. There are a lot of different options for catalytically avoiding the polymerization from forming at the top of the reactor. Some refiners include specialized iron trap catalysts which keep that from causing those formations.

LEICHTY (Chevron USA, Inc.)

Chevron has a number of units that store cracked feeds ahead of hydroprocessing units. When doing so, care must be taken to ensure that the tanks are free of oxygen to mitigate unwanted polymerization and gum formation. With the more reactive stocks like coker naphtha, it may be necessary to inject an antioxidant into the rundown and an antifoulant in the hydrotreater feed to mitigate exchanger fouling.

UNIDENTIFIED SPEAKER (Foster Wheeler)

Regarding the storage of cracked products, what is the maximum duration that we can normally store the cracked products? Can we take the cracked products from storage directly back into the hydroprocessing or should they go back to coker or somewhere else?

OHMES (KBC Advanced Technologies, Inc.)

As long as it is under a nitrogen blanket or floating roof, you should be able to store it for quite a while. I do not really have a good rule of thumb on storage time limits, but we have seen refiners go between turnarounds storing cracked stocks. But if you do not have that capability, the cracked products will quickly form polymers. You will then need to either try to rerun the material elsewhere or just take the material, knowing that it will cause problems in the units. Most refiners just get ready for it.

OHMES (KBC Advanced Technologies, Inc.)

Cracked stocks can come from conventional sources, such as coker, FCC, and visbreaker. However, cracked stocks can be found in unexpected streams, such as virgin stream where cracked refinery slop has been introduced to the upstream crude and vacuum unit. The hydroprocessing units most commonly impacted by cracked stock, from a fouling and pressure drop perspective, are the naphtha and diesel hydrotreaters. Some of the typical symptoms of fouling and pressure drop (DP) problems when running cracked stock are:

• Loss of heat recovery in the preheat train, such that the reactor charge furnace has to fire harder,

• Reduced hydraulic capacity, such that flow control valves operates further open or near full-open, and

• High pressure drop across the reactor or first bed, depending on indication points.

The normal cause of fouling and DP due to cracked stock is polymer formation, which is driven by the olefins and diolefins inherently in cracked streams, oxygen contamination in the feed, and catalyzed by iron contamination. The four management options are feed logistics, chemical treatment, unit configuration, and catalysts.

Feed Logistics

The primary cause of fouling from cracked stocks is O2 contamination. Therefore, all feeds to a hydrotreating unit processing cracked stocks should be sent from the source unit directly to the hydrotreater. If feed from tankage is sent to the hydrotreater, whether that feed is virgin or cracked, the tank should have a nitrogen blanket or should be a floating roof. Regular maintenance on the blanketing systems and floating roof seals are especially important for these services. Using the tanks in a “once-through” mode also assists in minimizing fouling.

As highlighted above, iron can serve as a catalyst to drive polymer formation. Some logistics mitigation options are to use feed filters to remove iron scale and to complete regular tank maintenance to remove corrosion products from streams coming from tankage.

Finally, ensuring that refinery cracked stocks are properly managed and processed is important. Several instances of high fouling have been traced to cracked slop being reprocessed in crude units, thereby “contaminating” virgin streams with cracked material. Normal practice is to segregate virgin and cracked slop and process cracked slop in the coker or send to fuel oil.

Chemical Treatment

Though chemical treatment is an option, the success rate is varied and depends on the severity of the fouling. For oxygen contamination, some refiners have used oxygen scavengers to contain the oxygen contamination. Other refiners have used polymerization inhibitors, filming antifoulants, and dispersant antifoulants to minimize the impact of fouling on the preheat exchangers. For reactors with pressure drop problems, chemicals are available for use to help break up the foulant and provide a temporary reprieve on the pressure drop. For these options, the best option is to discuss with your site chemical vendor to determine options and solicit positive result case studies or contacts.

Unit Configuration

Obviously, one option is to remove the cracked materials from the unit feed, but that option is normally not practical. However, an option may exist to process some or all of selected cracked streams in other hydrotreating units. For instance, some refiners process coker naphtha in the diesel or gas oil hydrotreater, as opposed to processing in a dedicated coker naphtha hydrotreater or mixing in the virgin naphtha hydrotreater, such that the material can be diluted in a larger feed pool.

Processing coker naphtha in conventional naphtha hydrotreating units is possible and common. However, the successful units include a diolefin saturation reactor. This reactor is loaded with low activity catalyst and operated at a relatively low temperature, as the diolefins do not require normal reactor temperature not normal activity catalyst to saturate. This configuration option allows the refiner to saturate the diolefins at low severity, thereby reducing the propensity to form polymers.

From a heat recovery standpoint, several options are available to minimize fouling impacts. First, ensuring some hydrogen is available in each of the preheat exchangers helps reduce polymer formation as the liquid feed is heated to reactor conditions. Most modern designs already have this feature, but older or revamped units can have liquid-only exchangers in the hot circuit that have a tendency to foul. Some licensors and refiners have installed the capability to take exchangers out of service while the unit is in operation to allow online cleaning or have sacrificial exchangers in the preheat train. Avoid using plate heat exchangers, which are prone to fouling. Finally, as a reminder for units which will operate in 100% vapor phase, make sure that the dry point does not occur in the charge furnace, so that deposits are not placed on the furnace tubes.

A final configuration options that some refiners have used is an oxygen stripper on the fresh feed. This option involves installing a low pressure stripper tower that uses hydrogen, natural gas, or fuel gas to strip out any oxygen contamination. However, this configuration option can be capital intensive and requires regular loss of off gas to fuel gas or flare to function properly.

Catalysts

Catalysts have long been the primary reactor pressure drop control mechanism. Most refiners now employ a catalyst grading system on the top bed of the reactor to provide a mechanism to trap contaminants and manage pressure drop escalation during a cycle. Specifically for fouling due to cracked stock, several considerations are suggested:

• In addition to catalyst size grading and progression, also ensure that catalyst activity is being progressed from low to high activity. This approach is similar to the discussions above on diolefin management.

• Special grading material is available that helps better trap and isolate iron scale. Doing so will reduce the tendency for iron catalyzed polymer formation. Discussions with your catalyst vendor about catalytic options should be part of your regular catalyst selection process.

ESTEBAN (Suncor Energy, Inc.)

Hydroprocessing of cracked feedstocks is very common due to the sulfur content of coker and FCC products. These feed streams can often be sent “hot” (directly) to a downstream unit, which prevents potential polymerization reactions due to oxygen ingress at storage tanks. However, storing these materials in feed tanks can result in reactor fouling and pressure drop increase.

Blanketing of tanks that store cracked feedstocks with nitrogen or some other gas is the primary control mechanism to avoid fouling and pressure drop problems resulting from the polymerization of cracked materials when oxygen is present. In some of Suncor’s facilities, cracked feedstocks are stored along with straight run feedstocks in nitrogen-blanketed tanks. In order to ensure proper operation the control instrumentation is inspected and maintained routinely to prevent fouling issues. In our experience, we have identified a malfunctioning N2 blanketing system as a contributing cause to exchanger fouling in a diesel unit. The malfunction had allowed oxygen to enter the tank and react with the cracked material in the tank.

A secondary means of preventing fouling issues is by the use of polymerization inhibitors available from chemical vendors. These are injected into the stream entering the feed tank. Most are approved for use, by licensors and catalyst manufacturers’. However, Suncor does not have experience with these products.

ROGER METZLER (Baker Hughes)

Avoid tankage, if possible. Direct hot rundown to the HDS unit. If intermediate tankage is necessary, nitrogen or fuel gas blanketing should be employed to minimize oxygen intrusion. Treatment with Baker Hughes antioxidants and/or polymerization inhibitors can be effective when the specific polymerization mechanism is identified in order to select the appropriate inhibitor chemistry.

BROOKS (BP Refining)

I would like to start with a little background on bottoms coke to present some clarity on the basis for my answer to this question. Bottoms coke sources could come from multiple places. But to address the question with respect to residence time, our impression was that we are mostly looking at the thermal cracking and decomposition reactions in the main fractionator bottoms as opposed to asphaltene precipitation or reactor transfer line coke generation and carryover.

BP has done a bit of internal research based on a 1988 paper by TOTAL that was presented in an AKZO catalyst seminar. We looked at the pool residence time, pool temperature, and asphaltene content. Based on that paper, we took those three pieces and combined them with some internal data to regress an equation. We then backed out a coking point temperature (based on about 0.45 wt% coke) as our coke formation point. We ended up with a coking point temperature, and that is what BP targets when we watch our main fractionator for bottoms coking. We do not focus as much on the residence time as we do on this coking point temperature. We have noticed that in this equation, when you look at the different variables of temperature, feed paraffinicity, and residence time, feed paraffinicity and temperature have much greater effects in our equation than residence time, which is why we focus on coke point temperature.

For most of our units, we use the same coke point temperature equation and track a certain temperature for the main fractionator bottoms pool. We will try to keep it about 10°F to 30°F below the calculated coke point temperature. That will then be one of our KPIs, or monitoring points, for our bottoms to help prevent coking in that circuit. Basically, to help control this, we focus on removing heat through the pumparound system. Also, 10 of our 11 BPoperated FCCs have an internal slurry quench system, which includes a slurry pumparound distribution ring near the top of the pool and a stripping steam ring near the bottom of the pool to provide extra agitation. Both of those rings will provide cooling and help us maintain the main fractionator bottoms’ pool temperature below the coking point temperature that I mentioned earlier.

To go back and address the question, though, since it was not really about temperature, BP calculates our residence times based on the main fractionator bottoms’ head volume, and then we divide it by the DCO product flow rate out of the unit. By that definition, BP’s residence times for most of our units are in the hour to hour and a half range. We feel that this is a fairly good proxy for the residence time in the circuit without having to go develop detailed calculations around the volume in the piping or in the exchangers.

If you are considering reducing the residence time, then it is conceivable for you to do this by dropping the liquid level in your main fractionator bottoms, which would then lower the residence time in your main fractionator. And if you are doing any major modifications, you can adjust the volume of the main fractionator head, although that is typically not cost-effective.

When looking at residence time as opposed to temperature, consider that as you drop the time in this pool, you will have a greater chance of building up coke into the bottom of the pool, which will result in extra clogging problems in the main fractionator bottoms circuit. You should always consider your pumparound rates and the minimum pump NPSH (net positive suction head) on your bottoms pumparound circuit. Lowering your residence time could reduce the amount of the NPSH available. But typically, BP focuses more on the temperature and less on the residence time.

PIMENTEL (CITGO Petroleum Corporation)

I have a little to add to what Halle said. CITGO’s basis for residence time calculation is the total volume from the normal fractionator operating level all the way to the inlet passes of the slurry feed heat exchanger divided by both the slurry and recycle. The sum of these two flows represents what is being purged out of the system. Again, there are not many options to reduce residence time without sacrificing the cycle oil yield than operating at your lowest possible fractionator level.

As for temperature control, we do find significant temperature variances between different points in the fractionator bottoms compartment after injecting the quench oil. So, I would recommend that you monitor temperature not only by the temperature of the product leaving the fractionator, but you should also take some readings around the bottoms to find the significant temperature variances.

WILLIAMS (KBR)

KBR designs for a one-minute holdup between the maximum and minimum levels based on the total slurry draw-off rate. Operating the tower at a lower level will reduce residence time within the column.

AVERY (Albemarle Corporation)

I agree with the comments made by Halle, Sergio, and Jesse. My detailed answer is in the Answer Book.

NAVEEN DIMRI (Reliance Industries Limited)

What is the typical CSO (clarified slurry oil) asphaltene content? Is there any limit for the fouling control?

BROOKS (BP Refining)

The equation we use does not include the asphaltene level specifically. The feed quality parameter is around the feed paraffinicity, which is obviously not really the asphaltene. We have that built in to account for feed quality. I have seen some other general references for the industry around feed quality. A lot of people will look at their minimum bottoms temperature and then subtract 20ºF or 25ºF off of that as an extra cushion around feed quality differences, if they have particularly nasty feed.

PIMENTEL (CITGO Petroleum Corporation)

We do not measure asphaltenes, but we do control the API. We have found that when we operate below -3 API or -4 API, the coking tendency increases significantly. So, we have set that as a limit for a minimum API of your slurry oil.

NAVEEN DIMRI (Reliance Industries Limited)

At BP, what is the asphaltene content you normally create?

BILL WILSON (BP Products North America Inc.)

Just to clarify, in this coking calculation, we actually looked at the paraffinicity of the feed. So that is the parameter. Normally, we do not actually measure the asphaltenes in the slurry or, in fact, even in the feed.

BROOKS (BP Refining)

Coke in the main fractionator bottoms section can originate from multiple sources. The key typical sources are coke formed in the reactor transfer lines, thermal cracking and decomposition in the main fractionator, and precipitation of asphaltenes. High catalyst fines in the bottoms can act as nucleation sites for the asphaltene precipitation and thus result in higher bottoms coke formation. Catalyst and coke entrainment can be considered as determined by factors outside the main fractionator. Thus, when focusing on controlling main fractionator bottoms coking, we focus on the key factors influencing thermal cracking and asphaltene precipitation: bottoms pool residence time, pool temperature, and asphaltene content. To avoid coking in the main fractionator bottoms section, we typically do not focus on the residence time as much as we focus on maintaining lower temperatures.

BP has done a significant amount of work internally around the factors that influence main fractionator bottoms coking. Based on a 1988 paper by Total at an AKZO catalyst seminar, the thermal decomposition of DCO is directly related to both residence time and pool temperature. BP used data from this study and additional internal data to regress a relationship between weight percent coke formation in the bottoms, residence time, and pool temperature. Additional studies followed which indicated that FCC feed paraffinicity also has a significant effect on these main fractionator coking reactions. Thus this factor was also added to our regression equation. For our purposes, BP assumes the critical coking point in the main fractionator bottoms at 0.45 wt% coke. The regression equation is solved for temperature at the 0.45 wt% coke level. This temperature is our “coking temperature.” The BP units regularly track the coking point temperature and work to maintain main fractionator pool temperatures 10°F to 30°F below this coking point temperature to prevent main fractionator bottoms coking. Typically, refiners have addressed this by specifying a maximum allowable bottoms temperature which is reduced by 20°F to 25°F for highly paraffinic feeds.

A holistic look at this equation suggests that the effects of temperature and feed paraffinicity on bottoms coking is much greater than the effect of residence time. Typically our main controls around temperature are to install slurry pool quench systems and focus on heat removal in the pump around circuit. Ten of our eleven FCCUs in the BP system have slurry pool quench systems to maintain low pool temperatures. These quench systems typically include a distribution system to deliver cooled slurry pumparound into the base liquid as evenly as possible. This stream should be submerged in the liquid, near the top of the pool. The system also typically includes a steam distributor near the base of the slurry pool to cool the bottom of the pool and aid in mixing the cooled slurry into the pool. This steam ring can have multiple benefits including:

- improving the stripping of the DCO and thus may also improve the DCO flashpoint,

- cooling the slurry in the base of the fractionator and thus reducing coking tendency, and

- providing agitation to help mix the quench and liquid in the fractionator base.

To focus on the base question around residence times, BP calculates the residence time by calculating the main fractionator bottom head volume based on liquid levels as the pool volume and dividing this by the slurry product rate. BP’s typical residence times are between one and 1.5 hours for this system. Some of our units run with residence times over two hours due to large main fractionator bottoms volumes and minimum bottoms make operations.

It is important to note that reducing the main fractionator bottoms volume in an effort to reduce residence time can have the potential to cause issues around build-up of coke shed from the transfer line. Without major equipment modifications, a site could conceivably reduce residence time by minimizing bottoms liquid levels and maximizing the bottoms pump around rates consistent with the bottoms heat removal required to maintain the desired fractionator heat balance. If major modifications or revamps are made to the tower (including a full tower replacement), and a goal is to reduce the residence time in the bottoms pool, modifications can be made to the shape of the bottom head of the fractionator. Changing the shape to a cone or other lower volume head will reduce the residence time. This is typically not a cost-effective modification in most situations which is why we focus on maintaining lower slurry pool temperatures (as mentioned before). If residence time is reduced, it is also important to review the pumps and revamp as necessary for lower NPSH required due to lower liquid head volumes.

PIMENTEL (CITGO Petroleum Corporation)

The basis for the residence time calculation are: the slurry oil flow (as produced from the unit plus flow recycled to riser) and the volume of the main fractionator bottoms compartment from the normal operating level plus the volume of the piping from the bottom of the tower to the inlet of the first set of heat exchangers. There are not many options to reduce residence time other than lower the main column bottoms level. We have found that maintaining a continuous flow of cold slurry quench into the bottom head of the main column reduces coke formation and subsequent plugging of the slurry pump screens or pumparound exchangers. This practice minimizes hot spots in the slurry liquid in the bottom of the tower where coke can form even if the bottom outlet temperature is at target (690°F max). If the bottom temperature drops below our target when the quench flow is at minimum, we reduce slurry pumparound duty at the steam generator. We have experienced slurry exchanger fouling due to the presence of excessive PNAs (polynuclear aromatics) and find that this often correlates with very low slurry API. This can be mitigated somewhat by sacrificing LCO recovery and dropping LCO into slurry.

AVERY (Albemarle Corporation)

Including the composition of the main column bottoms oil and the liquid temperature, an important variable impacting coking in the main column bottoms circuit is the hydrocarbon residence time in the bottom of the fractionator. Excessive time and temperature can lead to coke formation. The oil residence time can be determined from the volume of oil in the bottom of the tower and the bottoms flow rate leaving the tower. The residence time should be minimized by maintaining a minimum bottoms liquid level and maximizing the bottoms pumparound rate consistent with the bottoms heat removal requirement necessary to maintain the desired fractionator heat balance.

SUBHASH SINGHAL (Kuwait National Petroleum Company)

Lower fractionation, bottoms level, and optimum temperature (less than 365°C) are a few considerations to avoid slurry circuit fouling.

GATES (Motiva Enterprises LLC)

Again, we are trying to prevent the polymerization of the olefin/diolefin. The primary concern is trying to prevent contact with oxygen because that will ultimately lead to gum formation. So the preference would be, if possible, to feed this hot to all the downstream units and avoid intermediate storage. If you will have to put it into tankage, the preference would be to use an internal floating roof tank and add some type of chemical stabilization. So typically, some kind of oxygen scavenger would be put in there. The length of time that material will be staying in tankage may dictate the level of treatment that will be done.

The other concern is that if this material contains diolefin/olefins, as it goes into your feed/effluent exchangers, prior to entering the reactors, it could potentially foul those exchangers. Often, you want to add soak hydrogen or hydrogen upstream of that exchanger. It may need to be considerably upstream to ensure that as the material is going into the exchanger and getting hot, there will be some hydrogen present to prevent that gum formation.

If you are actually going to try and treat cracked naphtha, you will typically want a separate diolefin reactor that will be run at lower temperatures to avoid the polymerization. You will use catalytic hydrogen to eliminate the diolefin and operate it in a regime such that you will not saturate too many of the olefins. You are ultimately trying to put the naphtha into gasoline, so you want to minimize your octane loss for the gasoline pool. As far as the catalyst, look for a catalyst with sufficient surface area and large enough pore volume – going back to Sal’s comment – to manage any silicon or arsenic that might be coming with that cracked naphtha.

WATKINS [Advanced Refining Technologies (ART)]

Two of the reasons we are worried about olefins and cracked naphtha in our naphtha hydrotreater are that they are a large consumer of hydrogen and also because the reactions occur quickly. Once the reaction with olefins starts and we have enough temperature, it is hard to stop. We do not want excess temperature to coke up at the top of our hydrotreater. Like David said, with coker feeds, we would like to see them brought in straight from the coker while still hot. They should not be sent to tankage; and if they are, you should blanket the tank somehow and keep it protected from oxygen.

The top of the hydrotreater is where we are worried about excess exotherm. We want to spread out that heat and do the reaction somewhat slowly, if possible. We do not want to have it occur all at once. Generally, we will recommend a graded bed to mitigate that pressure drop and start olefin saturation to avoid worrying about anything else like HDS and HDN or aromatics. In that case, we will be using a low-metals guard material at the top of our hydrotreaters, which also gives us the ability to pick up silicon.

If your system has a high amount of diolefins, again, a dedicated diolefin reactor is important. At a much lower temperature, you can actually control just the diolefin reaction and avoid doing everything all at once by either controlling how much hydrogen goes to the hydrotreater or by keeping the temperature fairly low.

In the top of the main reactor of a coker naphtha unit, again, low metals, a guard material, is recommended. We generally then recommend avoiding the placement of very high-metals, high-active catalysts at the top, if at all possible, so we can try and do the reactions one at the time.

MUKESH PATEL (Reliance Industries Ltd.)

When we have to enable storage of naphtha, if necessary, which is the parameter we should monitor or what type of analysis should be carried out to determine if any gum formation has happened?

GATES (Motiva Enterprises LLC)

Off the top of my head, I do not know the ASTM (American Society for Testing and Materials) methods for that, but there are gums and then there are potential gums a

We run diolefin saturation tests, and I will put the ASTM method in my Answer Book response. We typically shoot for 60 to 80% diolefin saturation.

SHARPE (Flint Hills Resources, LP)

We run diolefin saturation tests, and I will put the ASTM method in my Answer Book response. We typically shoot for 60 to 80% diolefin saturation.

BRIAN WATKINS [Advanced Refining Technologies (ART)]

Processing coker naphtha can have several undesirable effects on the performance of the hydrotreater and the catalyst if the system was not properly designed to handle it. In general, coker stocks have a higher level of olefins present from the coking process. Once in the hydrotreater, these olefins will quickly get saturated (Figure 1), resulting in high hydrogen consumption and generation of a lot of heat. As a general rule of thumb, one mole of hydrogen is required per mole of carbon-carbon double bonds, or between five and 10 times the Bromine number reduction in standard cubic feet per barrel (scfb) of hydrogen. This additional heat [130 Btu/scf (British thermal unit per standard cubic foot) to 160 Btu/scf hydrogen consumed], if not managed properly, will initiate additional reactions, quickly creating a very high temperature rise. The high temperatures can accelerate coking and can lead to olefin polymerization, resulting in a dramatic increase in pressure drop. This can set an upper limit on how much coker naphtha can be processed based on the need to limit the heat rise or hydrogen consumption.

A system that is properly size and activity graded is extremely important when processing coker naphtha. ART recommends utilizing a grading system to help mitigate pressure drop buildup. A large inert hold down ring (GSK-19) with a very high void fraction used for trapping large particulates is placed at the top of the catalyst bed. A smaller diameter macroporous ring (GSK-9) that traps iron, as well as other fine particulates, is typically used in the next layer. After that, two types of smaller rings are used as active grading. These materials have a low level of active metals which help begin olefin saturation reactions, as well as provide additional particulate space at the top part of the catalyst bed. Avoiding the use of any highly active catalyst at this point is also recommended. Below the grading system, it is recommended to use a layer of larger size (1/10” or 1/12”) catalyst which provides activity for olefin saturation and additional void space for pressure drop mitigation. This layer is often a catalyst that is suitable for trapping silicon (and arsenic), which is another concern when processing coker naphtha.

An additional recommended practice to help prevent fouling and pressure drop buildup in coker naphtha units is to avoid contact of the coker feedstock with oxygen. It is preferred to bring the feed to the processing unit directly from the coker and avoid use of tankage. If this is not an option, then the alternative is to use a floating roof tank and purge the system to keep a nitrogen-blanket over the feed in order to keep oxygen out. Use of a diolefin treating reactor can also be considered. A diolefin reactor is operated at a much lower temperature in order to selectively catalyze the diolefin saturation reaction and avoid any excess heat generation from sulfur and nitrogen removal.

RAJESH SIVADSAN (UOP LLC, A Honeywell Company)

When processing cracked naphtha, plugging in feed exchangers and heaters may be caused by coking or polymer formation. Polymer formation can be caused by oxygen in the feed with cracked feedstocks. It is preferred to have cracked feeds fed to the hydroprocessing units directly from the upstream unit. Cracked feedstocks that do not come directly from the upstream unit should be stored in nitrogen-blanketed tankage. For units that process straight-run feeds from tankage, as well as cracked feeds, this straight-run feed should also be stored in nitrogen-blanketed tanks as this feed could provide the source of oxygen for polymer formation.

Processing these streams by themselves usually requires a two-reactor system with the first reactor operating at lower temperature to saturate the diolefins and a small portion of the olefins. The second reactor completes the olefin saturation and removes the sulfur and the nitrogen.

DENNIS HAYNES (Nalco Champion Energy Services)

Some issues may arise regarding polymerization. Polymerization in the feed stream may lead to fouling of heat transfer equipment and plugging of the reactor bed. A first step would be to minimize any potential for oxygen contamination as oxygen would act to accelerate fouling reactions. In addition, consideration may be given regarding the implementation of additive chemistries to inhibit polymerization rate (antioxidants) and minimize the potential for foulant (dispersants).

CHRIS CLAESEN (Nalco Champion Energy Services)

There are special low temperature reactor designs utilized to saturate the diolefins before they can thermally polymerize. Chemical programs with antipolymerants and dispersants can also help control polymerization and ∆P (delta P; pressure differential) increase. Reactor bed cleaning programs have been used successfully to reduce reactor ∆P caused by fouling with polymerized olefins and corrosion products.

RAJ PATEL (Haldor Topsøe, Inc.)

Designing a unit for processing coker naphtha can be challenging. The challenge comes from the silica in the feed which can poison the catalyst, high nitrogen in the feed which may be difficult to remove in low pressure units, and high quantity of olefin which can cause large temperature rises in the reactor, as well as very reactive molecules which are the conjugated diolefins. In the presence of small amounts of oxygen, or at elevated temperatures above 450°F, these molecules will radially polymerize to form gum that can foul exchangers or reactors causing poor heat transfer, as well as high rector pressure drop. If the feed contains a significant quantity of coker naphtha, then these diolefins must be removed at low temperatures to prevent gum formation.

The coker naphtha should preferentially be sent directly from the coking unit to the hydrotreater to prevent oxygen contamination. Even straight-run stock, which may be part of the feed component, must be prevented from contacting oxygen by storing the feed in a nitrogen-blanketed storage tank.

Even with strict adherence to avoiding feed contact with oxygen, the diolefins in the coker naphtha can polymerize at elevated temperatures. A dedicated diolefins reactor operating in the range of 300°F to 450°F will ensure that these highly reactive species are removed from the feed before polymerization can take place. Once the diolefins are removed from the feed, then the feed can be heated to the required temperature for silicon and/or sulfur/nitrogen removal. Topsøe has designed a large number of coker naphtha units with dedicated diolefin saturation units to mitigate the polymerization issues.

MIKE ROGERS (Criterion Catalysts & Technologies)

Hydrotreatment of cracked naphtha derived from FCCU, coker, or thermal cracking is a common refinery requirement. In some cases, refineries process these naphthas in a downstream catalytic reformer; while in others, hydrotreatment is necessary to meet gasoline blending sulfur requirements. Treatment of cracked naphtha can pose special challenges due to the presence of diolefins that are produced by the cracking reactions.

Diolefins in cracked naphtha streams from the FCCU and the coker will polymerize when heated, producing gums that foul the pre-heat equipment and main hydrotreater catalyst. To prevent polymerization, these naphtha streams are treated in a diolefin reactor operating at low but sufficient temperature to hydrogenate the diolefins. Typically, the diolefin reactor operates in a down-flow bubble phase with an LHSV of 2 hr-1 to 5 hr-1 and within temperatures between 180°C and 220°C (350°F to 420°F).

The effectiveness of diolefin saturation and protection against polymerization is monitored by measurement of the diene number in the reactor effluent. Typically, method UOP 326-08 (Diene Value by Maleic Anhydride Addition Reaction) is used to generate a MAV (maleic anhydride value) number representing approximately double the diene content in weight percent. Conversion of MAV in the diolefin reactor is in the range of 90 to 95%, and a target MAV number of less than 1 in the product stream from the diolefin reactor typically ensures protection against polymerization.

Catalysts loaded into a diolefin reactor should have a moderate but stable activity to selectively convert diolefins at low temperatures. For units processing coker naphtha containing silicon, a large surface area and high pore volume provide resistance to poisoning. Lastly, for units processing FCCU naphthas, catalysts with a low hydrogenation activity are favorable to minimize olefin saturation and octane loss.

JEFF CATON (Axens) 

Activation (or sulfiding) of hydrotreating and hydrocracking catalysts can be done either ex-situ or in-situ.  In both methods, the objective is to convert all the catalyst’s metal oxide sites to active metal sulfides (Refer to Figure 1 - Catalyst Sulfiding/Activation).  A poor or insufficient activation will result in reduced catalyst activity and cycle length.  This is applicable for nearly all hydrotreating and hydrocracking applications, including 1 and 2 stage hydrocrackers.  Axens’ recommendations do not significantly differ when considering activation (or sulfiding) a hydrocracker 2nd stage over a hydrotreater or hydrocracker 1st stage. 

For ex-situ catalyst activation (commonly referred to as pre-activation), the catalyst is fully activated via conversion of metal oxides to active metal sulfides.  Note, this is not the same as ex-situ pre-sulfiding, where the catalyst is more simply loaded with sulfur, but still requires an in-situ activation, or conversion of the metal oxide sites to active metal sulfides using the pre-loaded sulfur.  Pre-activated catalyst is delivered to the site in drop-in, ready to use form.  In addition to standard pre-activation, the catalyst can be additionally pre-treated to allow for immediate processing of cracked feedstock, and if applicable, can be pre-passivated in the case of hydrocracking catalysts. Additional treatment to allow for atmospheric loading of the catalyst is also available.  The disadvantage to pre-activation of catalysts is relatively limited to the cost of pre-activation, however, this cost is usually offset by the decreased loss-of-production time realized through elimination of in-situ activation (sulfiding) steps, which can commonly take 24-36 hours to complete.  Axens’ experience with the use of pre-activated catalysts has been exceptionally good, with most start-ups progressing quickly and smoothly followed by catalyst performing as expected.     

In terms of in-situ catalyst activation (sulfiding), the sulfiding can be done in liquid phase or in gas phase.  Further, the sulfiding can be done through injection of a sulfurous agent (e.g., DMDS, Sulfrzol, etc.) or via native feed/sour gas sulfiding, assuming a high enough sulfur content in the feed or sour gas.  All these methods are possible, and when completed according to recommended guidelines and best industry practices, with the proper safety and emergency scenarios considered, the catalyst activity and catalyst cycle length has achieved the expectations.   

Axens’ position is that liquid phase sulfiding is more controllable and robust, therefore less prone to unforeseen issues that might impact the overall catalyst activity and stability when compared with gas phase catalyst sulfiding.  The primary benefit of liquid phase sulfiding is the controllability of delta temperature (dT) caused by the exothermic sulfiding reactions.  Early in the in-situ sulfiding process, if the feed sulfur concentration is too high, the exotherm could result in a premature temperature increase creating a risk of coking or metal oxides reducing to the metal state.  If this were to occur, this could substantially impact the catalyst activity and stability, ultimately leading to poor yields and/or a shortened cycle length.  With liquid phase sulfiding, this risk is better managed as the liquid acts as a heat sink, thus dampening temperature spikes caused by the exothermic reactions. Similarly, Axens favors injection of a sulfiding agent versus native feed/sour gas due to the ability to better control the sulfur content going to the reactor through control of the sulfiding agent injection rate.  If the dT is too high and excessive hot spots occur, it is relatively manageable to decrease the sulfiding injection rate to decrease dT.  Conversely, with native feed/sour gas sulfiding, it is not as straightforward to monitor or control the amount of sulfur going to the reactor, so the dT is less manageable.  Further, based on Axens’ experience, sulfiding in liquid phase with injection of a sulfiding agent has generally led to shorter duration of the sulfiding procedure. 

A risk during in-situ sulfiding is the loss of sulfur to the reactor caused by loss of sulfiding agent injection.  If there is a loss of sulfiding agent injection during the higher temperature step, there could be a resultant reduction of metal sulfides back to metal oxides and then further to the metal state.  With liquid phase sulfiding, it is relatively more manageable to cool the reactor temperature to a lower temperature where the risk of reduction is minimized.  With gas phase sulfiding, it is more difficult to decrease temperature of the reactor in a timely manner.   As a precaution to counter the risk of loss of sulfiding agent injection, Axens recommends that the oil and/or makeup hydrogen used for sulfiding contain a minimal amount of sulfur and that the recycle gas amine scrubber be idled throughout the sulfiding process.  These precautions will ensure continuous sulfur levels in the system thus minimizing the risk of catalyst reduction.  Axens recommends that the recycle gas H2S content not fall below 1,000 vppm H2S when the reactor temperature is above 400 oF. 

Like the ability to cooldown the reactor, it is also found to be more manageable to heat-up with liquid phase sulfiding.  This can save time during the initial heat-up of the reactor to the first temperature hold and subsequently to the second temperature hold.  Even when the refiner is limited to complete sulfiding in gas phase, Axens would recommend circulation of treated diesel throughout the unit to ensure wetting of the catalyst pores and to provide heat sink as discussed above. 

It is important to note that in-situ liquid sulfiding should be completed with a 100% virgin (straight run) feedstock.  After full completion of in-situ sulfiding, it is recommended to operate the unit with 100% straight run feedstock for a period of 3 days and then gradually introduce cracked feed stock.  Further, the initial operating temperature should be maintained below the expected SOR temperature and then slowly increased as required to meet product specifications.  Operating the unit with cracked feedstock within the first 3 days, increasing temperature too rapidly, and over treating may result in premature catalyst coking and resultant loss of activity and cycle length.  

NEIL HOWARD (Chevron Lummus Global) 

For current generation catalyst CLG has experienced comparable catalyst performance with either gas phase or liquid phase sulfiding procedures.  There are some units that require gas-phase sulfiding, mostly because of the mechanical configuration – high MPT reactors, inadequate heat input, inadequate recycle compressors, etc.  Chevron’s experience suggests that gas-phase sulfiding can be done safely and timely and without causing any yield differences, even in a second-stage reactor.  We have observed that some older catalysts give slightly worse start-of-run performance under gas-phase sulfiding, but many of the most modern catalysts show little or no reduction in activity with gas-phase sulfiding. 

Liquid phase vs. Gas phase startup 

In general liquid phase sulfiding is preferred for reasons already stated above.  Some of these liquid phase advantages include: 

• better heat sink during the sulfiding steps to manage temperature rise in the catalyst beds. 

• faster response to temperature increases or decreases in the event of a need to suspend sulfiding activities in the course of the startup. 

• faster heat-up to MPT on units where startup oil is introduced at low pressure under nitrogen. 

• no prolonged cool-down to feed-in temperature after high temperature sulfiding 

Some of the concerns with heat release with gas phase sulfiding mentioned above can be mitigated by: 

• Maximizing recycle gas rate to improve distribution and carry heat from the catalyst beds during sulfiding 

• Limiting initial DMDS injection to maintain H2S in the recycle gas to no higher than 0.5 vol%.  The reactor temperature rise should be limited to 50 degF 

• DMDS injection can be increased to 1 vol% H2S in recycled gas after the initial temperature wave passes through the reactor.  But again, dT should be limited to 50 degF maximum. 

• sulfiding can be conducted at low pressure to avoid delays due to heating up to MPT and going to full pressure.  400 psia minimum H2 partial pressure is recommended, or as high as allowed by MPT. 

MAUREEN BRICKER and JOHN KULACH (UOP) 

This question requires some background on catalyst manufacturing and studies that we have performed utilizing both activation procedures on various hydrocracking catalyst types.  As catalyst performance is being optimized to enhance yield and targeted hydrogen saturation so too are the methods employed for catalyst formulation and manufacture.  Each step of the process is studied and optimized for the exploitation of the catalyst performance. Optimization techniques in manufacture and the activation of the catalyst in-situ are linked to catalyst performance.  While the gas phase activation tends to maximize the metal sulfide uptake there is a clear advantage with liquid sulfiding pertaining to maximizing the yield and in many cases the activity of the catalyst system.  During liquid phase sulfiding, there is some modest coke formation even at mild liquid sulfiding conditions.  The type of coke formation occurring during the activation phase likely prepares the catalyst better for initial design feed by preventing initial over activity and higher rates of coke formation during early stages of the cycle.  It could also attenuate the most acidic sites on the catalyst which would deactivate immediately in process feeds. 

Liquid sulfiding procedures may be more difficult to enact in older units where the minimum pressurization temperature (MPT) may be higher than the potential catalyst reduction temperature in hydrogen atmosphere. Since liquid phase sulfiding is performed under hydrogen atmosphere, there is concern that heating the catalyst under hydrogen to achieve minimum pressurization temperature (MPT) could result in catalyst reduction. This is not a concern with gas phase catalyst activation. UOP has developed procedures for those reactors requiring higher minimum pressurization temperatures to enable use of liquid phase catalyst activation.  

Regarding temperature excursions: In gas phase and liquid phase sulfding, the sulfding reaction results in axial rise across each bed of catalyst.  In liquid phase sulfding exotherms can be seen additionally from adsorption (wetting) and from cracking.  Procedures must take these possibilities into account and provide avoidance or action to be taken.   

BULL (Valero Energy Corporation)

I will initially address this question from a CO standpoint and then discuss the NOx. CO emissions typically increase if the regenerator falls below a certain temperature threshold. That temperature threshold will vary based on your regenerator configuration and definitely on the type of air distribution. In general, well-mixed regenerators with longer residence times can be run at significantly lower temperatures before CO emissions increase, compared to a poorly mixed regenerator with lower residence time.

In the absence of a CO combustion promoter, large variations in the CO2 (carbon dioxide)-to-CO ratios are observed. At the catalyst surface, it is believed that that CO2-to-CO ratio is an intrinsic function of the temperature at the burning site. However, the CO exiting the burning site may be further oxidized to CO2 at a rate dependent on the temperature, amount of CO and O2, water partial pressure, and active metals on the catalyst.

To mitigate CO emissions at low regenerator temperature, we add a non-platinum CO promoter to limit the NOx. Other solutions include increasing 1) your delta coke, 2) the pressure in the regenerator, or 3) the mixing by modifying the regenerator. The amount of NOx emitted from the regenerator is highly dependent upon several variables such as operation regenerator style, excess O2, promoter concentration co-distribution, feed nitrogen, and dense bed temperatures. So, there are many variables that determine the overall NOx equation. NOx formation chemistry inside the regenerator is difficult to pin down. Studies have shown that about half of the feed nitrogen becomes coke on the catalyst, but only a small portion of that actually becomes NOx. It is also believed that NOx emissions are limited by the reaction of CO plus NO (nitric oxide) to form N2 (nitrogen gas) and CO2. Thus, when you have platinum CO promoters added to the unit, CO is converted to CO2 so quickly that there is less CO available for that reaction to occur.

We do have an experience of the regenerator being in true countercurrent flow. With platinum promoter in the unit, it is able to run NOx levels that are very low; but, I will say that is a rarity. In our system, if there is normally a platinum promoter, then you will see elevated NOx levels. In partial-burn operation, there is always CO available to react with the NOx; so emissions are lower. Good coke and air distribution is important to ensure that the concentration of CO, O2, and NOx are evenly distributed throughout the regenerator. The results have been better and more efficient control of the burn zone so that oxygen is used more efficiently and the regenerator can be operated at lower levels of excess oxygen to completely burn the coke to CO2. Generally, you do not need to use platinum promoter. Finally, uniform residence time distribution minimizes the duration that spent catalyst remains the regenerator.

Due to the interaction between CO, NOx, and the FCC regenerator, I recommend that you actually conduct testing to determine the tipping point where you start going up in CO emissions because there are benefits from lowering the oxygen at that time. Running O2 in excess slightly above that tipping point provides a good NOx baseline of your minimum level on the unit. If that level is still too high, then you have other options which include LoTOx™, WGS+™, CONOx™, and SCR (selective catalytic reduction). We have all of these applications across our system. We have very little experience with NOx reduction additives.

GIM (Technip Stone & Webster)

The low regen temperature used to be beneficial. However, with the advent of tight oil and light feed stock, actually too much of a good thing is now bad. So let us first consider what causes some of these low regen temperature phenomena.

The first cause is a result of better feed stock qualities, such as lower Conradson carbon, lower nitrogen, and metals. Second, lower operating severities, such as lower reactor outlet temperature and lower catalyst activity, will reduce regenerator temperatures. The third cause is catalyst regeneration, including full- versus partial-burn, CO promoters (platinum-based or non-platinum-based), and supplemental oxygen that has been turned off.

Jeff described many of these aspects, but I thought it would be worthwhile to summarize all of the factors into three main categories. One is, obviously, the combustion kinetics. Second is the hardware of the actual generator design; and third, to a certain aspect, are the feed contaminants such as feed nitrogen. In terms of combustion kinetics, a lot of these items have been discussed multiple times in past panels, so I will just go over the basics and summarize them. The lower regenerator temperature increases the CO but decreases the NOx or NO2; the promoter lowers the CO. But in terms of the distinction between a platinum and non-platinum base, the promoters can be either increased NO2 or NO2 neutral. The higher O2 concentration lowers the CO and increases the NOx. Again, these are some of the factors that you can use to counteract the formation of CO and NO2.

BART de GRAAF (Johnson Matthey INTERCAT, Inc.)

All of the speakers mentioned that the low temperature does affect the CO oxidation and the formation of NOx. The nitrogen kinetics in the regenerator are not affected by the lower temperature. Nitrogen in coke can be converted into nitrogen in NOx, hydrocyanide, and ammonia. As was suggested, non-platinum combustion promoter can be used to steer the oxidation reactions from HCN and NH3 towards N2 instead of NOx. Partial-burn operation is one example where you can see that the CO concentration does affect high NOx. But at the same time, in a partial-burn operation, after the CO boiler, you frequently see a higher NOx level than you would have in a full-burn operation because of the high temperature oxidation in CO boiler.

KEN BRUNO (Albemarle Corporation)

I want to point out that Albemarle has a detailed answer in the Answer Book. But in summary, Albemarle’s non-platinum promoter, ElimiNOx™, provides high CO promotion, excellent afterburn, and minimal impact on NOx emissions.

ROBERT “BOB” LUDOLPH [Shell Global Solutions (US) Inc.]

If you are an oxygen enrichment operator or considering oxygen enrichment, this is your opportunity to raise your bed temperature at the same coke burning rate and shift the CO and NOx distribution within the bed. In turn, the control of NOx and CO emissions could improve.

MICHAEL WARDINSKY (Phillips 66)

We have done a lot of modeling around NOx and CO emissions for additive use. We have seen NOx emissions decrease with higher regen temperatures in almost all of the modeling cases. From our experience, we also know that the non-platinum promoters can increase NOx emission, although not as severely as a platinum-based promoter. We have moved partial-burn units from deep partial-burn to full-burn and observed the NOx emissions decrease substantially. I think it was from the gentleman from INTERCAT who mentioned that in partial-burn, you are releasing ammonia and HCN (hydrogen cyanide) off the regenerator in greater quantities than during full-burn operations. However, these reduced nitrogen species are combusted across the CO boiler to NOx, so your partial-burn units always appear to emit higher NOx emissions than operating in full-burn, even after accounting for thermal NOx generated in the CO boiler.

JEFFREY BULL (Valero Energy Corporation)

CO emissions will increase if the regenerator falls below a certain temperature threshold. That temperature threshold varies based on the regenerator configuration and the type of air distribution. In general, well mixed regenerators with longer residence times can be run at significantly lower temperatures before CO emissions increase compared to a poorly mixed regenerator with lower residence times. In the absence of a CO combustor promoter, large variations in CO2/CO ratios are observed. At the catalyst surface, it is believed that the ratio of CO2/CO is an intrinsic function of the temperature at the burning site (“Arthur's ratio”). However, the CO exiting the burning site may be further oxidized to CO2 at a rate dependent on temperature; CO, O2, and H2O partial pressures; active metals on the catalyst; carbon/oxygen distributions within the fluidized bed; and even the catalyst presence.

To mitigate CO emissions at low regenerator temperatures you can:

• Add non-Pt (platinum) CO promoter to limit NOx,

• Modify your operation to increase your delta coke (catalyst or operational changes),

• Increase pressure in the regenerator, and/or

• Modify the regenerator to increase mixing.

The amount of NOx emitted from the regenerator is highly dependent upon several variables such as mode of operation, regenerator style, excess O2, promoter concentration, coke distribution, feed N2, and dense bed temperature. Studies have shown that about half of the feed N2 goes to coke on catalyst, but only a small portion of the N2 in coke goes to NOx. It is believed that the NOx emission is limited by the reaction of CO + NO to form N2 and CO2. Thus, when Pt CO promoter is added to a unit, CO is converted to CO2 so quickly that there is less CO available to react with the NOx; so NOx emissions increase. We do have experience with a regenerator that is in true countercurrent flow, and the operator is able to run at low NOx levels with platinum promoter in the unit. In a partial-burn operation, there is always CO available to react with the NOx formed; so, emissions are lower. Good coke and air distribution are important so that concentrations of CO, O2, and NOx are evenly distributed throughout the regenerator.

The results of the better, more efficient control of the burn zone in a regenerator are that:

• Oxygen is used more efficiently so that lower levels of excess oxygen are required to completely burn the coke to CO2 while minimizing CO emissions.

• Generally, no Pt combustion promoter is needed (at design rates) to accelerate the burn to completion.

• The uniform residence time distribution permits the time spent in the regenerator to be minimized.

Due to the interaction between CO and NOx in an FCC regenerator, testing should be conducted for your individual regenerator to determine the “tipping point” where a reduction in excess O2 causes CO emissions to increase dramatically. Running at an excess O2 slightly higher than the “tipping point” provides a baseline NOx that you can expect for your unit. If this level is too high, then you will either need to modify your regenerator for better mixing or use an alternate method of NOx reduction like LoTOx™, WGS+™, CONOx™, or an SCR. We have four LoTOx™ applications, one WGS+, one CONOx™, and one SCR in our system. In some cases, a catalyst additive can reduce NOx, but we have limited experience with this application.

STEVE GIM (Technip Stone & Webster)

Why lower regenerator temperatures? Let us first examine some of the factors leading to lower regenerator temperatures: first, better feedstock qualities (such as lower Conradson carbon), lower nitrogen, and metals. Second, lower operating severities, such as lower reactor outlet temperature and lower catalyst activity, will lower regenerator temperatures. Third, catalyst regeneration, including full versus partial-burn, CO promoters (platinum-based or non-platinum-based), and supplemental oxygen turned off are additional options.

Factors Directly Affecting CO and NOx: Many of these directly affect the formation of both CO and NOx in the regenerator. Some of these factors influence the formation of these two species in opposite direction. I have summarized these factors into three categories and their directional changes of the CO and NOx. Many of these items are obvious and have been extensively discussed in the past panels, but I thought it would be a good summary.

Interactions Between CO and NOx: Reductant-like CO converts NOx to elemental nitrogen in the presence of O2. Lower regenerator temperature can be the direct result of higher CO to CO2 ratio in partial-burn units, which in turn will lower the NOx.

CHRIS STEVES (Norton Engineering)

The impact of low regenerator temperature on CO and NOx emissions can be very unit-specific, depending especially on regenerator design and the distribution of air and coke in the regenerator. In full-burn units, operation at low regenerator temperature will generally result in higher CO emissions and lower NOx emissions. The use of CO promoter can help in reducing CO formation, but use of platinum-based promoters is well documented as increasing NOx emissions. Non-Pt-based promoters are widely used today to allow for control of CO (and afterburning associated with CO) while limiting NOx emissions.

RYAN NICKELL (Albemarle Corporation)

Though not directly stated, this question is most applicable to FCC units concerned with after-burn. Lowering the regenerator temperature can indeed lead to an increase in CO and NO to N2 and CO2. A platinum-based combustion promoter can be used to reduce CO emissions with excellent success. However, NO will increase as intermediates, such as HCN and NH3 (ammonia), are also oxidized by platinum. As a result, many refiners prefer a non-Pt-based promoter such as Albemarle’s ElimiNOx™. This additive provides high CO promotion activity but acts less on the aforementioned intermediates containing nitrogen. ElimiNOx™ provides excellent after-burn control and minimal impact on NOx emissions.