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CARLSON (Criterion Catalysts & Technologies)

Catalyst typically deactivates due to an accumulation of contaminants and due to process upset coke formation. It also deactivates due to what I will refer to as routine coke formation. Routine coke formation is two-dimensional, its rate being a function of endpoint, as well as the feed aromaticity and how much polyaromatics you have in it. The coke that we generate from the high endpoint tends to come from cracking, both catalytic and thermal, hopefully not too much thermal in our units, whereas the coke from the aromatic content comes from condensation reactions.

The example on this slide is a bit of an eye chart. It is an example of a noise analysis of an LVGO (light vacuum gas oil) feed. The only reason I have it up there is to demonstrate the two-dimensionality of the feed, endpoint versus its hydrogen, and level of hydrogen saturation. The gray line on the chart show the little arrows going into the coke precursors where the deactivation rate becomes too hard for typical operation. It is somewhat arbitrary as it is also strongly a function of unit partial pressure, space velocity, and your operating temperatures. I put it here just to illustrate the fact that from a rate of coke perspective, straight run feeds, of course, can have much deeper endpoints than the cracked stock feeds with a high amount of PNAs in them.

In a typical operation when our customers are also adhering to reviewing the level of other contaminants, we have good, long operational cycles for VGO (vacuum gas oil) straight run feeds with endpoints of 1050°F and 850ºF for 100% FCC cycle oil. Again, these are arbitrary, depending on the unit configuration, and can be optimized as we are going through our operating cycles. So that is where we would start.

When pushing endpoints, you should carefully monitor the result. For a single-feed component, as the endpoint is increased, it will bring with it more of the difficult refractory material that has a higher propensity for coke formation. It is the “tail” of the distillation that contains most of the undesirable coke precursors. The impact of this “tail” can be managed by setting limits on the maximum 95% or 98% boiling point. Other important feed measurements to monitor include metals (Ni, V, Fe, As, Si, Na, etc.), Conradson carbon, nitrogen, and asphaltenes contents.

Catalyst deactivation is not the only guiding consideration. This figure illustrates how the polyaromatics can grow into larger polycyclic aromatics, or PCAs, such as coronenes or ovalenes that we run into during hydrocracker operations.

In a high conversion hydrocracker employing a recycle stream, PCAs can accumulate and grow into heavy polycyclic aromatics (HPCAs) due to a combination of the high residence time provided by recycle and the low reactivity of these species. HPCAs are orange-to-red (hence, the name “red death”); and in high concentrations, they cause increased catalyst deactivation, deterioration of yield selectivity, and REAC (reactor effluent air cooler) fouling. FCC cycle oils contain high concentrations of HPCA precursor; and when the feed endpoint is increased, more of these HPCA precursors will enter the unit. “Red death” can be managed with the appropriate amount of unconverted oil purge and/or endpoint control. For optimization, start with something that is reasonable, push until you see your unit’s capability gets a little challenged, and then pull back from there.

OHMES (KBC Advanced Technologies, Inc.)

I want to add a couple of more points to build onto what Kevin said about really understanding how you are trying to use your hydrocracker. Is it processing all available feed, or are you putting in the really difficult material? What is your ultimate distribution for the incremental cutpoint material? For instance, on the coker, yes, everyone wants to run at minimum recycle, but maybe that is not the best option if the heavy coker is going to the hydrocracker. The same with the gas oil: Are those last few degrees of cutpoint really worthwhile? Because it is really the last 5°F, 10°F, or 15°F of cutpoint that are going to dictate the severity, both in the hydrotreating and the hydrocracking section of the hydrocracker. So, trying to optimize not only the hydrocracker itself but how it fits into the overall refinery is the key.

With that as a background, we typically recommend our spec set not only be the blended feed, but you should also look at the individual feeds. So, the old rule of thumb that people would say was, “2 ppm nickel plus vanadium; keeping it less than that and less than 0.5% Conradson carbon.” Again, to Kevin’s comment, those are good starting points, but you need to optimize around that, as well try to understand the distribution in your feeds. I know this is hard to see from the back of the room; but depending on where you are running conventional or unconventional crudes, you can see what those distributions look like. These are generated from Petro-SIM, which we used to help understand these distributions. So, when you are looking at that cutpoint, a simulation tool really helps you figure out what the optimization needs to be. We typically recommend test runs, as well as some model work and proactive catalyst cycle life management, to figure out what that endpoint needs to be.

CARLSON (Criterion Catalysts & Technologies)

One other comment to mention is the importance of the modeling. The rigorousness of the modeling is useful as well because an LP (linear program) might show more benefit for going beyond 1050ºF than there really is. You will see that when you put in a barrel of that next cutpoint material, it really does not do anything for your liquid yield gain. So as Robert mentioned, reviewing how you are modeling the impact of that is quite important.

OHMES (KBC Advanced Technologies, Inc.)

One other quick point to the earlier question HPNA (heavy polynuclear aromatics) management: Again, the same principle applies. You have to ask yourself if you really want to go after that last one percent conversion, or would it be better to just bleed that out and maybe give gain the capability of heavying up the feed without plugging everything up with HPNAs. That is an operability and economic decision that each hydrocracking refiner has to make.

KEVIN CARLSON and WARD KOESTER (Criterion Catalysts & Technologies)

When also adhering to other guidelines (limits on trace contaminants and CCR), our customers have achieved long cycles in high-conversion hydrocrackers by limiting the endpoints of straight run VGO to 1050°F and 850ºF for 100% FCC cycle oil. However, the true limit for a given unit and a given feed depends on other factors, such as (a) process configuration, (b) catalyst type, (c) hydrogen partial pressure, and (d) treat gas rate.

When pushing endpoint, you should carefully monitor the impact on unit performance. Just as importantly, you should be sure that your economic models accurately represent the value of higher endpoint. Molecules that boil above 1050ºF produce far more coke and far less liquid than lighter molecules. It is commonly observed that adding a barrel of high-endpoint material gives zero (or negative) incremental yield of converted product.

For a single feed component, as endpoint is increased it will bring with it more of the difficult, refractory material that has a higher propensity for coke formation. It is the “tail” of the distillation that contains most of the undesirable coke precursors and the impact of this “tail” can be managed by setting limits on the maximum 95% or 98% boiling point or the amount of 700+°F or 1000+°F material in the feed. Other important feed measurements to monitor include metals (Ni, V, Fe, As, Si, Na, etc.), Conradson carbon, nitrogen, asphaltenes/heptane insolubles, and aromaticity/paraffinicity.

With a mixed feed, the combined endpoint does not always provide the best indication of that feed’s propensity to coke. That is because from the standpoint of deactivation, you can operate to higher endpoints with straight run material than you can with poly-ring material. This is due to the relationship between endpoint and molecular structure on deactivation.

Figure 1 shows the H2 deficiency (Z value) of the compounds in a straight run LVGO versus the carbon number (N) for those compounds. The boiling points listed on the Y-axis are for paraffins. Paraffins have H/C ratios of (2N+2)/N and Z = +2. Forming a one-ring naphthene removes two H atoms, so Z = 0. Forming a benzene ring removes six more H atoms, so Z = -6. For coronene, a large polyaromatic compound, the formula is C24H12, the H/C ratio is 0.5, and Z = -38. The numbers in each cell represent the wt% of compounds with a given Z and N; components with Z = -12 and N = 16 to 18 are the most concentrated. -As shown by the coke-precursor curve across the middle of the chart, for a given value of N, compounds with low Z (such a polyaromatics) are more likely to form coke than compounds with high Z (such as paraffins).

Figure 2 illustrates how polyaromatics can grow to form larger polycyclic aromatics (PCAs), such as coronene and ovalene

In a high conversion hydrocracker employing a recycle stream, PCAs can accumulate and grow into heavy poly cyclic aromatics (HPCAs) due to a combination of the high residence time provided by recycle and the low reactivity of these species. HPCAs are orange-to-red (hence the name “red death”) and in high concentrations they cause increased catalyst deactivation, deterioration of yield selectivity, and REAC fouling. FCC cycle oils contain high concentrations of HPCA precursor, and when the feed endpoint is increased, more of these HPCA precursors enter the unit. “Red death” can be managed with the appropriate amount of unconverted oil purge and/or endpoint control. You can determine the appropriate amount of purge required by monitoring the amount of coronene and ovalene species in the unconverted oil. The rules of thumb are to keep the coronenes at less than 100 ppm and keep the ovalenes less than 10 ppm.

With unprotected pretreat catalysts, our customers have achieved long cycles with straight run conventional VGOs having endpoints of 1050°F, and with 100% FCC cycle oil feeds having endpoints of 860°F. When the main catalyst is protected with appropriate grading material, the endpoints for straight run feeds can be higher.

Color is an excellent indication of feed quality. Red FCC cycle oils are likely to cause problems. Black VGO is also bad. For such feeds, dilution is not a solution. In one case, mixing just 500 bpd (barrels per day) of black oil into 30,000 bpd of hydrocracker feed did not change the SimDist (simulated distillation) endpoint, but it caused a loss of 40°F of activity in less than two months. In another case, raising the endpoint of an HCO feed from less than 860°F to greater than 900°F doubled the deactivation rate.

LP planning models may not account properly for endpoint effects. A model is wrong if it estimates that a barrel of resid (>1100°F) gives the same yields as a barrel of 650°F to 1050°F material, most likely, that barrel of resid will produce considerable amounts of coke and very little 650°F minus.

OHMES (KBC Advanced Technologies, Inc.)

To supplement the responses given by the other panelists, the outside batter limits (OSBL) implications and impacts should be examined. First, the refiner should clarify how the hydrocracker is being used within the facility. Though the question relates to distillate mode, that mode can be achieved through conventional feedstocks (virgin gas oil and coker gas oil), but may also involve processing other feedstocks (Deasphalted oil, FCC LCO, coker naphtha, etc.). In addition, is the hydrocracker processing relatively “easy” feeds, such as light vacuum gas oil, or “hard” feeds, such as full range heavy coker gas oil?

Also, setting the right endpoint targets depends on the overall conversion level. For instance, if the unit is targeting greater than 95% total conversion, processing high back-end distillation feed can have a dramatic impact on hydrotreating and hydrocracking reactor temperatures. However, if a higher unconverted oil rate is acceptable or even economically desirable, a heavier endpoint can be tolerated. Some good examples of high value unconverted oil dispositions are as FCC feed or Lube Complex feedstock.

Many hydrocrackers operate in multiple modes, depending on unit configuration, unit flexibility, and seasonal economics. Therefore, if the hydrocracker operates in a distillate mode on gas oil feeds and a naphtha mode on distillate feeds, the needs of both modes impact the amount of heavy feed that can be processed. For instance, if the hydrotreating catalyst has “spare” activity and the unit is going into distillate mode on gas oil feed prior to turnaround, the back-end quality targets can be altered to fully consume the catalyst prior to change-out.

Finally, cycle length targets impact the decision on where the hydrocracker feed back-end quality specifications should be set. Depending on marginal cutpoint and stream routing economics, operating the unit at a different cycle length than the standard two years may be economics.

To adjust the back-end specification on hydrocracker feed, an upstream unit cutpoint must be adjusted or stream routing changed. Therefore, to select the right specification, the alternate routing of these streams or cutpoint material must be included. For instance, if the heavy vacuum gas oil cutpoint is lowered to reduce the contaminant load on the hydrocracker treating catalyst, this cutpoint material will now become fuel oil or coker feed, which impacts refinery economics and operability. Adjusting coker heavy gas oil cutpoint affects not only the fractionator’s performance, but also the coker recycle ratio and throughput. Therefore, the hydrocracker backend specifications have to be set in the context of the overall facilities economics and reliable operating targets.

As stated by others, the hydrocracker back-end specifications should be set by contaminants, reliable and reproducible distillation points (i.e., 95%) and methods (simulated distillation or D1160), and impact to polynuclear aromatic (PNA) formation. The refiner should monitor not only blended feed qualities, but also individual feed stream qualities. Monitoring only the blended feed can cause critical contaminants or poor fractionation to be masked or hidden. For example, many hydrocracker units monitor Conradson carbon in the last 10% of the distillation on streams to properly understand the Conradson carbon content, as monitoring Conradson carbon on the bulk stream can dilute the content.

Understanding the distribution of contaminants across the boiling range will influence how the back-end distillation target is set for a given feedstock. The following charts provide some typical examples of how Conradson carbon and vanadium are distributed for different benchmark crudes, based on KBC’s Petro-SIM simulation tool.

The recommended practice for setting back-end specifications for hydrocracker feed is to use a combination of predictive tools and simulations, such as Linear Program and kinetic model within a process simulator, along with proactive unit monitoring. The tools can help optimize the stream routings and cutpoints to achieve a refinery profitability and operability target, whereas unit monitoring can validate the predicted performance and provide feedback into the analysis, as well as allow for fine tuning of targets against evolving facility and unit limitations. The bottom line is that increasing the back-end distillation of hydrocracker feed will negatively impact the unit operation and cycle length, but the economic benefit must outweigh the penalty. Technical and economic analysis on the specific unit, refinery configuration, and market condition are required to make the proper decision.

SUBHASH SINGHAL (Kuwait National Petroleum Company)

The endpoint of the feed is fixed by the cycle length that the refiner intends to achieve. The higher the endpoint, the more coke on catalyst and the more PNA in feed to the second-stage which may result in exchangers fouling up and catalyst deactivating in the second stage. I have seen HCU operating in diesel mode with feed endpoint of 1020°F and cycle length of two years.

LEICHTY (Chevron USA, Inc.)

There have been many inquiries into shifting naphtha hydrocrackers to distillate mode over the last couple of years. The optimal solution is dependent on feed qualities, hydrocracker configuration, distillation, and gas recovery hardware, as well as the value of volume expansion across the hydrocracker.

When it comes to feed qualities, higher distillate yield can be achieved by shifting away from paraffinic feeds to more naphthenic feeds because of the improved cold properties. Another way to enhance distillate yield in a hydrocracker is to feed stocks that are rich in diesel boiling range material, such as FCC LCO. LCO boils almost entirely in the distillate range and contains around 80% aromatics, which translates into high volume expansion when saturated with hydrogen. Minimizing feed overlap will also increase distillate yield against a cold flow property constraint since this material has higher amounts of n-paraffins than the material derived from cracking.

The strategy for maximizing distillate will also depend on hydrocracker configuration. For a single-stage hydrocracking unit, option one would be to maximize the more-difficult-toprocess cracked feeds already in the diesel boiling range and minimize hydrocracking to lighter products. Option two involves switching to a heavy VGO feed and then to a maximum distillate and cold flow property selective catalyst. In either case, it is best to configure the reactor temperature profiles to be flat or only moderately ascending in order to minimize over-cracking.

The next slide illustrates the second option and shows how switching from a more naphtha selective catalyst like ICR183 to ICR 185 in a single-stage hydrocracker not only gives you the cold property improvement shown in Question 8, but also results in improved distillate yield.

Now let’s look at a two-stage hydrocracker with intermediate distillation. The same strategy of increasing distillate yield by feeding cracked or otherwise-difficult-to-process diesel boiling range materials can have additional benefit. In addition to the high first-stage yield we observed with a single-stage unit, these materials also help a two-stage unit make more distillate because they can reduce the per-pass conversion required in the second-stage reactors. Lower per-pass conversion results in higher second-stage distillate selectivity because there is less potential for over-cracking. If the two-stage hydrocracker has tail-end distillation, flat second-stage temperature profiles, coupled with a distillate selective catalyst, provide the best opportunity to increase distillate yield.

Whether the two-stage unit is intermediate distillation or tail-end distillation, there are some new improvements that can be made to the second-stage catalyst. The next slide shows the difference between two newer base-metal catalysts relative to the previous generation distillate selective catalyst: ICR 240. With ICR 250, you can choose to make more diesel at the expense of gasoline, or you can employ ICR 255 for a more jet-selective catalyst with a significant drop in start-of-run temperature.

When it comes to distillation and recovery, the principles are the same as those we discussed for improving cold flow properties. Maximizing internal reflux prevents heavy tails from impacting cold properties, and good stripping of the bottom's product minimizes distillate lost to unconverted oil. On the front end of the distillation, dropping lighter product down to a distillate flash limit will provide some cold property carry and allow additional heavy molecules to be pulled, thereby maximizing the width of the cut. If the fractionator only has a jet draw, it may be possible to increase the recycle cutpoint of your jet draw, and then route it to diesel. Doing this would require changing to a diesel flash specification but may result in more overall distillate. Another option would be to add a separate diesel draw and side stripper to the hydrocracker main fractionator. This option could require upgraded metallurgy in the bottom system because the temperatures will go up. Alternatively, a vacuum column could be added to remove diesel from the unconverted oil.

Last, but most important, to consider when switching from maximum naphtha mode to maximum distillate mode is the value of volume expansion. Shifting the hydrocracker to distillate mode will result in a loss of volume expansion. Since products are generally sold on a volume basis, the loss of volume can overwhelm the price difference between distillate and gasoline, especially if the hydrogen being added to the oil is derived from inexpensive natural gas, as is the case in North America.

BODOLUS (CVR Energy)

Steve did a good job covering a lot of the issues and concerns. As he mentioned, the concerns are similar to those of Question 7, but they come at it from the other higher conversion end. We have a list of procedures to follow when trying to improve the distillate slate with diesel being over gasoline, but I think probably the biggest recommendation is to rationalize the changes in the naphtha production, the impact it has on the hydrogen balance at your refinery, and also how competitive it will be to switch to diesel, in terms of FCC economics.

CARLSON (Criterion Catalysts & Technologies)

You two covered it really well, so I will minimize my response. Most of the conversion capacity of the approximately 60 operating hydrocrackers in North America is geared to the gasoline market. Nevertheless, many still have the flexibility of switching from gasoline mode to distillate mode (by adjusting process operating parameters such as cracking conversion, liquid recycle rate, and product cut points) without requiring capital expense.

Just as a reminder, this topic is fully discussed in NPRA Paper AM-09-10, “Dieselization in North America: Flexible Solutions for Diesel Production” by Robert Karlin.

Operational adjustment changes: Shifting to a distillate mode operation is typically achieved by switching from a recycle operation to a once-through operation while lowering conversion, which increases diesel production. When reducing conversion to increase diesel yield, a lower overall conversion level may result in lower distillate quality due an increase in aromatics and resultant drop in API gravity.

A drop in product quality can be overcome through the use of a more distillate selective catalyst that will improve distillate quality via higher hydrogenation activity while improving distillate yields.

Lowering conversion also results in reduced hydrogen consumption. Increasing feed rate to utilize the available hydrogen can recover the losses in volume gain across the hydrocracker. Therefore, one primary modification to gasoline hydrocrackers for the improvement of distillate yield and throughput involves revamping the fractionation section to enable higher feed rates.

BODOLUS (CVR Energy)

Considerations start with feed slate and catalyst selection. For shifting to distillate production, the feed slate will need to be heavier to produce a higher yield of distillate range materials. Catalysts are also specific to feed and product distribution. Catalyst vendors can offer yield projections for various feed slates to match desired product distributions. It is important to get assurances from the catalyst vendor that produced diesel will meet local specification for sulfur, distillation and cold flow properties.

Once the feed slate has been selected and product yields have been determined, hydrogen consumption and distribution within the reactor need to be evaluated to assure compressors and control valves are sized to deliver the required quench streams. Downstream distillation equipment and product rundown may also have to be modified to accommodate the changes in product slate. Key assurances need to be obtained on the distillate product qualities to make sure they will be compatible with prevailing specifications on sulfur, cetane, and cold flow properties.

Shifting hydrocracker selectivity to distillate production has wide ranging implications to other units in the refinery. Overall changes in economics need to consider if the hydrocracker is competing for FCC feed and if the drop in naphtha production impacts hydrogen supply from the reformer operation.

CARLSON (Criterion Catalysts & Technologies)

As global diesel demand has risen and is forecasted to continue its growth, North American refiners have been compelled to evaluate their processing options to expand diesel production. Most of the conversion capacity of the approximately 60 operating hydrocrackers in North America is geared to the gasoline market. Nevertheless, Shell Global Solutions has helped refiners discover their flexibility for switching from gasoline mode to distillate mode by adjusting process operating parameters such as cracking conversion, liquid recycle rate, and product cutpoints without requiring CAPEX (capital expense).

Further increases of diesel production and quality from existing assets can be achieved through changes in catalyst to a distillate selective system from Criterion/Zeolyst, as well as through unit revamp projects. Clearly, each hydrocracker is unique and requires a detailed analysis of feed diet, operating constraints, and desired yield to adjust and optimize operations. Maintaining flexibility to adjust to diesel market conditions and building in the capability to handle a variety of feeds will allow for increased hydrocracker profitability.

Hydrocracker distillate mode operation is typically achieved by switching from a recycle operation to once-through operation while lowering conversion, which increases diesel (bottoms) production. When reducing conversion in an HCU to increase diesel yield, a lower overall conversion level may result in lower distillate quality due an increase in aromatics and resultant drop in API gravity. A reduction in product quality can be overcome through the use of a more distillate selective catalyst that will improve distillate quality via higher hydrogenation activity while improving distillate yields.

Lowering conversion to increase distillate yield will also result in reduced hydrogen consumption. Increasing feed rate to utilize the available hydrogen can recover the losses in volume gain across the hydrocracker. Therefore, one primary modification to gasoline hydrocrackers to improve distillate yield and throughput involves revamping the fractionation section to enable higher feed rates.

Hydrocracker revamp scenarios for maximizing distillate range can range significantly in scope. Shell Global Solutions has successfully increased distillate yields through minor modifications of the fractionation section of hydrocrackers. In gasoline hydrocrackers where light and heavy naphtha are drawn from the main fractionator, the heavy naphtha draw has been redesigned as a distillate draw. This modification increased the cutpoint of the fractionator bottoms and increased the boiling range of the cycle oil, considerably reducing the cracking of light and middle distillate products to light naphtha. Another common hydrocracker revamp involves upgrading reactor internals. Shell Global Solutions has experienced some gasoline hydrocrackers designed for an all-vapor phase quench zone requiring new internals based on changes in feed and conversion. It is well documented in North America and throughout the world, that this minimal capital investment strategy can increase liquid yield. These benefits have been seen in over 300 hydroprocessing applications in hydrocrackers owned and operated by refiners world-wide.

For some refineries, the opportunity to bring heavier feeds into a hydrocracking unit is desirable; but in cases where the fractionator bottoms are diesel pool blending material, adding heavier feed to the hydrocracker is likely to render the diesel unsuitable for blending. The addition of a vacuum flasher to recover diesel uncouples feed back-end from diesel cold properties and enables the processing of heavier/better feeds for increased distillate yield.

In larger scope revamp scenarios, significant increases in unit feed rates can be achieved thru the paralleling the stages (converting to single stage). In one North American refinery where this type of revamp occurred, a portion of the FCC feed was directed to the revamped hydrocracker in return for high quality bottoms product from the hydrocracker. A number of other modifications are required for this type of revamp to succeed, including the fractionation section.

SUBHASH SINGHAL (Kuwait National Petroleum Company)

Units designed for naphtha can be used for distillate based on shift in product demand and resulting financial benefits. Catalyst systems need to be changed for distillate selective yields (lower activity catalyst). A change in reactor severity will also provide enhanced distillate yields.

LEICHTY (Chevron USA, Inc.)

The answer to this question primarily centers around the type of reactor, but there are some additional considerations.

With an ebullating-bed reactor, the primary throughput limit is set by the superficial velocity, which is a function of reactor diameter and throughput. The superficial velocity is what determines the ability to separate gas from liquid going to the recycle pump. The largest diameter reactor in service today is 13.5 feet ID (inside diameter), which allows for a feed rate of 47 kbpd (thousand barrels per pound) to 50 kbpd. It is now possible to build a 15-foot inner diameter that would allow the feed rate to increase between 65,000 bpd and 85,000 bpd, depending on the conversion level desired.

With a fixed-bed reactor, the throughput limit is set by the desired reactor flux which, again, is a function of reactor diameter and throughput. The next step is to calculate pressure drop across the reactor and each piece of equipment to ensure a good design. It is also important to consider the ability to distribute the reactants in a very wide bed and to uniformly mix quench hydrogen for temperature control. The ASME 2000 code does provide for higher allowable stress, which makes larger reactor diameters possible. With that change in code and CLG’s latest ISOMIX internals, a 70,000-bpd hydrocracker or 85,000 bpd hydrotreater can be designed in a single-train unit.

The reactor diameter and throughputs discussed thus far assume that there are no other limitations or restrictions. There are other considerations that could reduce the single-train capacity. For instance, if a refinery is landlocked, there may be diameter and/or weight limitations set by roads, port infrastructure, or bridges. Additionally, not every manufacturing shop is capable of fabricating to the maximum diameter. There may also be some site-specific erection limitations.

Aside from the reactor, there are other practical limits that may prevent reaching the aforementioned throughputs. Pressure drop through the high pressure loop can increase to the point where a recycle compressor cannot deliver the desired capacity. If that is the case, care must be taken during the design due to the fact that heat exchanger tube bundle diameters max out at 60 inches and that the maximum furnace tube diameter is about 10 inches as a result of heat flux. CLG prefers to design with two-pass furnaces, but four-pass has been done and would allow a reduction in pressure drop. High-pressure loop-piping diameter limitations can also become a significant source of loop pressure drop.

Finally, it is important to think about the desired cycle length, onstream factor, and future debottlenecking capability. With a single-train design, there is no ability to take down one module at a time for a catalyst change while increasing rate on the sister module. It may also be harder to manage oil movements during catalyst changes with a single-train than with a dualtrain design. If future debottlenecking is anticipated, then it might be easier to go with multiple modules rather than a single reactor given that a single-train design may be stretched closer to all of the aforementioned limits.

OHMES (KBC Advanced Technologies, Inc.)

I want to add a couple points to what Steve said. In the mid-2000s when everyone was trying to build all of these new units, it was not just the reactors that were causing problems for the CHPSs (cold high-pressure separators). Often, you could not find the pumps and compressors of a desired size to put into these units, which would somewhat dictate whether you did multiple-train units or went one of two or two of three in your makeup compressors.

We primarily focus on ensuring that our clients understand what will happen during a future shutdown; because early in the FEL (front-end loading) process, someone wants to minimize capital. So yes, that may be the right decision, but you will need to put in a kit that will function in the refinery. So, if that unit is down and you have all your eggs in that basket, then what will happen to that feed? Do you have to shut down the whole plant? Do you have to change your crude slate or whatever it might be? Do you have the ability to sell that feed, or are you landlocked and cannot do so? Are you going to be able to meet product specs? In our cases, we are often able to justify two smaller parallel trains in order to keep one online to sustain the entire facility operational. I will not go through these in a lot of detail because there is more in the Answer Book.

We have worked with a few clients to do some Monte Carlo scenario analysis to help them understand the implications of one versus two trains. Also, the ebullated bed units typically, unfortunately, have a slightly lower on-steam factor than do some of the conventional hydrotreating technologies. So, you need to understand the implications to the plant and what the units can achieve. You also need to get down to the equipment level to really understand that if you recycle or makeup compressors are going to be more problematic, then you will want to focus more attention in those areas during the design phase. In summary, you want to balance not only the capital needs, but also the overall operating constraints to the facility so that when you do start-up and run this, you will be able to achieve the targets you wanted.

BODOLUS (CVR Energy)

We have already licensed a unit of 70,000 bpd unit, single-train, and fixed-bed reactors. It will have a single makeup gas compressor because it has a centralized hydrogen compression system. So that is not considered a bottleneck. What I do understand is that the constraint to limit the feed to 70,000 barrels was max flux rate. I am not able to understand how this it is related to the superficial velocity; so, if you have any answers to that situation, please respond.

OHMES (KBC Advanced Technologies, Inc.)

We agree with Steve’s points. I know there are some in Asia that are in the 70,000 bpd to 75,000 bpd range. So it is possible.

SUBHASH SINGHAL (Kuwait National Petroleum Company)

Reactors’ size and weight may limit the capacity of a single-train reactor within maximum mass flux constraints.

OHMES, DAVID WILLIAMS, and MIKE VOIGT (KBC Advanced Technologies, Inc.)

We agree with the stated size limitations and considerations made by the other panelists. However, several other aspects outside of reactor design come into the decision-making process. For instance, the ability to secure pumps and compressors (recycle and makeup) becomes a more difficult challenge as the unit’s size increases. Though the normal convention is to have a single recycle compressor with two makeup compressors with some multiple of expected capacity, large units may need to consider more makeup compressors at lower capacity per machine to be able to have the machines constructed in the time required for the project. Similarly, large multi-stage pumps can be difficult to secure during busy construction periods.

In the early stages of the front-end loading (FEL) process, capital is a major constraint and consideration. Therefore, most refiners strive for the configuration and unit size that minimizes capital outlay. However, as the project progresses through the FEL process, the practical operation of the unit and facility comes into play. Therefore, KBC normally stresses to clients that the impact of single versus multiple units/trains should be included early the FEL process, particularly on how the facility will operate during catalyst changeouts and turnarounds. For instance, if the hydrocracker will have a catalyst changeout every two years and full turnaround every four years, where will the hydrocracker feed go for that 15- to 45-day period each time? Will the feedstock be sold, stored in tankage for later processing, or will facility throughput have to be cut back or even severely curtailed? Depending on the facility location and market conditions, the lost opportunity cost for these situations may drive the refiner to consider two trains or two units to mitigate the lost revenue. In addition, how will the loss of the unit’s product affect product blending? For instance, can the facility still produce jet and diesel products without the high-quality hydrocracker streams? An analysis of additional capital versus lost opportunity savings should be completed early in the process and communicated to the respective licensors.

Equipment and unit reliability have a similar impact on this decision. Therefore, the effective onstream factor should be generated to determine how the overall facility will operate when respective equipment is down for regular and unanticipated maintenance. For instance, ebullated bed hydrocrackers have significantly improved in onstream factor in recent years. However, by comparison to most modern refining technologies, those units struggle to consistently achieve greater than 90% onstream factor. Therefore, for roughly one month per year, the ebullated bed unit will be out of service. This outage will have a significant impact on the overall facility and may drive the decision to reduce the size of the ebullated bed unit. The ebullating-bed process requires evaluation and selection of the optimal downstream processing scheme to maximize yield and product quality at the minimum investment cost. The investment ultimately will depend on the capacity and utilization of the ebullating-bed unit in conjunction with the capacity and utilization of the existing units to process the products and residual oil from the ebullating-bed unit. Options for the residual oil include for example: coking and gasification. Typically, additional hydrotreating of the ebullating-bed products is required for both sulfur and hydrogen content. The optimal design for the complex system can be determined with the application of reliability modeling using the Monte Carlo method.

As an example, KBC completed a detailed system analysis as part of a major capital expansion project for a complex refinery. Through stochastic reliability modeling, multiple scenarios were evaluated based on various system configurations, unit process technologies, unit sizing and turn-up capacities, equipment and process failures, and turnaround strategies, to determine the optimal process design.

From the system analysis, individual unit performance, overall system performance, and sensitivities for each scenario were appraised. To illustrate this approach simplistically, a naphtha hydrotreater (NHT) is examined. This particular scenario has a modified NHT unit with turn-up capacity to 105% of design. The following graph represents the crude unit utilized production capacity over a mission time of 10 years for this scenario. The two major interruptions are the result of a plant-wide shutdown on a four-year cycle. In this scenario, the average utilized capacity for the crude unit is 91.4%.

The production distribution over this same 10-year mission time is presented in the following graph. The crude production utilization distribution indicates that the performance of the crude unit is between 89% and 93% with a 50% probability of exceeding 91% of design capacity.

For this example, the top ten major events that contributed to production loss for this complex refinery are represented in the graph for Production Loss Sensitivity. The plant-wide turnaround (TAR) is the major contributor to production loss accounting for 14.2% of the total losses. TAR strategies and alternatives for sequencing can then be evaluated to reduce this sensitivity and quantify overall system performance improvements. Other high probability production loss scenarios can be reviewed in a similar fashion to determine how design or configuration changes will impact overall availability and onstream factor.

Finally, facility plot space and infrastructure have an impact on unit sizing. Though the facility may want a large single-train unit, the plot space may not be available for the unit and the associated support units (hydrogen plant, utilities, etc.). Emergency flare size can also impact this decision, as larger units require larger flare systems. Based on the relief scenario analysis and reliability analysis, the size of the flare lines and flare itself may make a large unit difficult to build.

Therefore, in addition to examining the issues associated with designing and constructing the reactors for large single-train units, the refiner should also consider the implications of unit size on pump and compressor sizing and construction, managing planned and unplanned outages, equipment reliability, and overall facility plot space and infrastructure needs. Ultimately, the refiner should strike the right balance between capital outlay and operating costs and constraints

ROBERTSON (AFPM)

We have three questions coming up that the panel chose to answer only in the Answer Book to keep the question count down during the session. Question 14 is one of them. Kevin has a very detailed answer in the Answer Book, but we are not going to present the answer here. The panelists felt strongly about including these three questions in the Answer Book.

ANDREA BATTISTON (Albemarle Corporation)

Factors limiting cycle life in low pressure kero/jet hydrotreaters are mainly related to product quality specifications. In particular, specifications related to aromatic saturation such as smoke point and color.

Jet has very tight specifications for color and color precursors and is required to pass a specific test, the JFTOT (Jet Fuel Thermal Oxidation Test). The JFTOT measures color of oxidized deposits in a kero/air mixture at 527°F. Color detected by the JEFTOT is believed to be caused by precursors very similar to those responsible for color seen also with other color measurements, although the JFTOT can often reveal such species at lower concentrations [as low as 1 wppm.

The main constraints for kero hydrotreating operation are typically HDS and HDA activity. HDS is relevant, in particular, for ULSK production, for blending with on-road diesel (to improve cold diesel properties), city diesel (to lower soot formation) and sometimes with off-spec ULSD to reduce blended S content. HDA is mainly relevant for improving smoke point, but care must be taken in feedstock selection for low pressure kerosene units since the capability to saturate aromatics is limited.

In some instances, pressure build-up can be a limiting factor too, in particular when operation is easy and long cycles are achievable from a catalyst activity standpoint. A well-designed guard bed system with activity and catalyst shape grading for trapping of scale can help mitigate pressure drop issues.

Limitations in kero hydrotreating for both HDS and HDA become more severe when hydrogen pressure in the unit is low and when feed severity increases. In both cases SOR WABT is higher as a result, together with catalyst deactivation rate. As a consequence, a higher temperature is reached earlier in the cycle (as a rule of thumb, a 10°F increase of the feed T95 above 475°F corresponds to a 5°F SOR WABT rise due to the larger presence of refractory sulfur and nitrogen species, and of aromatics). Temperature becomes a serious problem in operation when feed becomes near fully vaporized. Heavier species left on the catalyst surface are more susceptible to dehydrogenation reactions. Traces of 2+ ring aromatics may cause color. In some cases, it may be desirable to raise temperatures sufficiently to fully vaporize the feed rather than operate near, but not quite, full vaporization.

Proper catalyst choice is important for achieving maximum cycle length in kero/jet low pressure operation while meeting product specifications. In general, CoMo catalysts with moderate hydrogenation activity, like Albemarle’s KF 757, are preferred at low pressure when HDS is the main target because of their higher HDS activity and robustness with low hydrogen availability. A combination of CoMo/NiMo catalysts can also be applied to increase HDN and HDA activity, for instance, to promote smoke point improvement.

Units with small operating windows may benefit from even higher HDS activity, like Albemarle’s KF 770. In low pressure commercial operation, KF 770 has demonstrated SOR WABT can be decreased by ~10°F compared to alternatives, leading to significant cycle length increase by delaying onset of temperature related color problems.

GREG ROSINSKI and BRIAN WATKINS (Advanced Refining Technologies)

For most jet and kerosene hydrotreaters, the end of run condition limitation is typically the product color or smoke point. Aromatics and poly-nuclear aromatic (PNA) compounds are a problem for both smoke point and color. Color bodies are PNAs that form as the reactor temperature is increased. Several of these types of molecules are green/blue and fluorescent in color, and the color is apparent at very low concentrations of these species. Certain nitrogen (and other polar) compounds have also been implicated as problems for distillate product color. These species can polymerize to form various color bodies and can also form sediment via oxidation and free radical reactions. At end of run conditions when the outlet temperatures are high enough, PNA saturation reactions become equilibrium limited, the formation of aromatic compounds is more favorable.

If the unit is having difficulties with color, several steps can be taken to extend the cycle length. Increasing the H2/oil ratio to the hydrotreater will increase the reactor outlet H2 partial pressure helping to slow down the reverse equilibrium reaction and can help in prolonging cycle life. Feeds such as light cycle oils and coker stocks typically have higher H2 consumption than straight run materials, which reduce the outlet H2 partial pressure. Increasing the available hydrogen to these units will help in maintaining the kinetic reaction. If unable to add additional hydrogen, minimizing cracked stocks will also help with color.

If the unit has enough quench available, operating with a descending temperature profile for the last bed can also assist with increasing cycle life. This also has the added benefit of improving the outlet hydrogen partial pressure which as mentioned above is beneficial to the PNA saturation reactions.

MUTHU SRINIVASA (Criterion Catalysts & Technologies)

The cycle limiting factors in low pressure ultra-low sulfur jet/kerosene (ULSK) operations are often tied to meeting color specification for jet fuel and limiting the vaporization fraction at reactor outlet to avoid dry point for ULSK (<10 ppm sulfur). ULSK EOR temperatures are generally higher than those of jet fuel production.

Factors that influence jet product color include both process conditions and feed composition.

Process Conditions:

• High outlet temperature

• Low H2 partial pressure (pressure, % H2, treat gas rate)

• Low LHSV

• Product sulfur (over treating well below 10 ppm)

• High reactor outlet vaporization reducing H2 pp (partial pressure)

• Bad distribution (localized hot spots)

Generally, low H2 partial pressure and less than 10 ppm product sulfur require reactor outlet temperatures beyond the aromatics equilibrium point resulting in both high concentration of diaromatics and poor color quality. Maximum reactor outlet temperatures allowable to control color vary and is a function of feed quality, H2 partial pressure and LHSV. Some refiners operate the unit alternatively between jet mode and ULSK with limited tankage flexibility. This requires operating jet also to less than 10 ppm sulfur with a flexibility to blend with ULSK. This high severity operation can lead to color limited cycles.

Feed Composition:

• High aromatics (low API)

• Feed high T90/T95 distillation

• Feed source determines the feed relative processing difficulty. It has been observed that the feed color is much lower for Kerosene derived from high asphaltic or synthetic crudes as compared to sweet crude kerosene.

• High feed nitrogen (>50 ppm to 100 ppm) leads to higher product nitrogen beyond aromatics equilibrium temperature and results in bad color.

ULSK production is often blended to ULSD and hence color is not typically a major issue. EOR conditions for ULSK in low pressure hydrotreaters can be determined by maximum allowable temperature to avoid dry point in reactor. Low pressure and lower H2 partial pressure requires higher reactor outlet temperature to meet less than 10 ppm which leads to higher vaporization. It is preferable to restrict reactor outlet vaporization to have better control on reaction kinetics

ROBERTSON (AFPM)

Since we only have about 40 minutes left to get through 12 questions, we are going to just take primary answers now. The secondary and third answers are in the Answer Book.

OHMES (KBC Advanced Technologies, Inc.)

We are starting to hear about more foaming problems, particularly with some of the unconventional or shale crudes being processed, and there are a few other instances. First of all, I will admit that a lot of these were covered in the early 1990s in the NPRA Q&A, and I shamelessly borrowed from them, noting that I had done so. Also, there was a really excellent article by Turner et al from Marathon that discussed the problems they had and how they went about to solving them. Obviously, we have run into a few of these issues as we have talked with clients.

So just briefly, refiners are typically finding some amount of foaming occurring. One will normally see either some liquid carryover, problems with level control where you are not truly seeing the level itself, and/or a lot of hydrogen under carry. I will skip most of these slides since I already touched on them.

Ultimately, the problem is really around separation. Maybe you have expanded the unit and did not take a look at the separator, thinking it would be okay when, in actuality, it did not have the capability for good separation. There are a few other problems people have seen, particularly if they have a recycle hydrocracker with a lot of HPNA, such that foaming will occur. As an example, Marathon has units processing feeds with high asphaltenes, particularly those with DAO (deasphalted oil). Another minor problem we have heard is that several people changed the process chemistry of their washwater and were getting some chemicals that were causing some foaming. Exchanger leaks can cause this, obviously, if you are running higher temperatures.

What are some options that you can use to fix this problem? First of all, we would start looking at your plant data. If you see a lot of hydrogen under carry showing up downstream in your separators, fractionator, or stripper, you can compare that quantity to what the thermodynamics will tell you to determine if there is an opportunity to not only stop the foaming, but maybe also recover some hydrogen. You could do some sampling to check the leaks. But honestly, what most people have had to do to isolate the problem is just do some test runs, pull out feed to check for a hydraulic issue, and play with some of the individual feeds, whether they are DAO or not.

So a few of the mitigation options: There are chemicals available, and we have seen people use them. To be honest, it is a bit hit-and-miss. For the ones who have been successful in using chemicals, their use is normally a short-term gap to get them through to a shutdown in order to fix the separator. From most of what we have seen, people have either had to replace or revamp the separator.

One example occurred when we worked with a company by the name of EGS. They design and manufacture Vortex Tube Clusters, which are a variant of a hydroclone. Although it is a little hard to see on the diagram, these tubes basically sit down within the liquid, which forces the vapors to go through the tubes. By going up and through the tubes, the emulsion is broken, thereby giving you the benefit of not only minimizing the foaming but also reducing the hydrogen undercarry. So with this particular client, we provided some of the process data because we were doing unit operability improvement work; EGS developed the revamped design. The client saw such a significant reduction in the gas losses that they were able to increase severity and throughput, thereby reliving the hydrogen constraint. So to summarize: Look at your plant data, and do some test runs. But for all intents and purposes, you are probably going to have to do some type of a modification to your separation equipment.

CARLSON (Criterion Catalysts & Technologies)

This phenomenon has tended to be observed with hydrotreater multistage recovery section designs when a change in the hot high-pressure separator (HHPS) operating condition led to a foaming/level control problems in the cold high-pressure separator (CHPS). When observed, the foaming in the CHPS appears to be very sensitive to the temperature at which the HHPS is operated at with foaming indicated above a particular range. This is most likely due to a change in the vapor liquid equilibrium or possible entrainment resulting in heavier than expected material being routed to the CHPS resulting in a tighter HC/H2O emulsion forming after injection of the washwater. The higher HHPS operating temperatures leading to this problem may have been due to operating the reactors at a higher EOR temperature or from fouling in the F/E (feed/effluent) exchangers. This issue can often be mitigated by the cleaning of F/E exchanger and somewhat managed by a modified descending quenching strategy of the catalyst beds.

OHMES (KBC Advanced Technologies, Inc.) and DAVID LEAKE (EGS Systems, Inc.)

These particular questions were addressed in the early to mid-1990s NPRA Q&A sessions, specifically 1991, 1994, and 1997. In addition, an article was published in 1999 by Turner, et al that addressed the issue of foaming and emulsions (Hydrocarbon Processing, June 1999). The following response includes the information from those sources, as well as KBC’s experience in dealing with foaming and emulsion problems in hydroprocessing units.

Separator designs assume gravity separation and residence time will be sufficient to separate gas and liquid. The following is a brief description of the design process:

• Density between gas and liquid favors the separation.

• Interaction between the liquid and gas forms a stable fluid or emulsion/foam.

• Gas is entrained with the liquid and leaves in a higher concentration than pure component solubility will predict.

• The result is the separation that does not match design and recoverable hydrogen leaves with the liquid.

Causes and Indicators of Foaming and Emulsions:

Based on these information sources and KBC’s own experience in this area, the following are the causes of foaming and/or emulsions in high pressure separators:

• Hot Separators

o Inadequate separation capability increases with throughput or recycle gas rate,

o High content of polynuclear aromatics (PNAs) in hydrocracking seems to form stable emulsions, and

o High asphaltenes in feed, particularly when processing deasphalted oil (DAO) or resid.

• Cold Separators

o Chemicals in injected washwater,

o Exchanger leaks, and

o High temperatures in separator.

The typical indicators that a unit is experiencing foaming or emulsion problems are:

• Hot Separators

o Higher than expected hydrogen and light ends losses,

o Liquid carryover from separator into downstream exchangers resulting in loss of heat transfer, and

o Level control instability.

• Cold Separators

o Level control instability,

o Significant water in fractionation section downstream of separators, and

o High recycle compressor vibration.

Mitigation Options:

Prior to selecting a mitigation method, the refiner determines the root cause of the problem or at least narrows down the list of potential causes. The first indication is that the fractionator or stripper off gas hydrogen content is much higher than design. This situation also can cause the fuel gas amine absorber to become over-loaded. Recycle gas flow is less than design and heat is lost in the feed/gas exchangers leading to over-firing of the charge heaters.

To determine if hydrogen and light ends soluble losses are high, the reconciled mass balance and elemental balances (carbon/hydrogen/sulfur/nitrogen) are compared to a simulation of the unit’s light ends and hydrogen loss. If the actual losses are significantly higher than the value based on thermodynamics, then the separator likely has a problem, and the potential improvement value can be calculated. This same approach, when accompanied by a kinetic model and some select sampling of additional unit streams, can help validate if exchanger leaks are occurring.

For validating potential causes such as operating conditions or feedstock type, a series of short-term test runs can be completed. If the unit is processing a new high asphaltene content feed, then removing the feedstock may reduce the foaming. Of course, replacing the feedstock to maintain relatively constant unit throughput will allow verification of feedstock type versus throughput as a cause. Upstream fouling of exchangers may be leading to elevated separator temperature, such that the remedy is exchanger cleaning rather than a unit modification. Finally, washwater quality and source should be validated to ensure that unknown chemicals are not being injected into the water and driving the foaming or emulsions.

Chemical addition to improve foaming or break emulsions are an option. However, based on the past NPRA responses and KBC’s experience, the success rate is “hit or miss”. Those units that have had success are able to properly disperse the chemicals in the separator to achieve the desired improvement, so some type of an injection system will be needed. A short test run with chemical injection may give a preliminary indication of improvement. Most users have avoided chemicals as a long-term solution due to the typical costs.

The most successful mitigation of foaming and emulsions is to improve the separator through revamp or replacement. KBC has direct experience with such a modification. For a client who was experiencing significant separation problems in a resid hydroprocessing unit, KBC partnered with EGS Systems, Inc. to revamp a high pressure separator, with KBC providing the necessary process data and EGS providing the design. The primary revamp involved installation of vortex tube clusters (inlet defoamer) internals. The advantage of these designs is no change is made to the high pressure containment vessel. The cyclones are internally supported, and no attachment welds are required.

The following schematics, courtesy of EGS Systems, Inc., provide generic examples of these vortex tube clusters installed in a horizontal and vertical separator.

For this particular installation, the revamp greatly improved separation, such that the entrained gas losses decreased dramatically and unit hydrogen utilization increased. Therefore, the economics payback was a matter of months, not including the impact of foaming and emulsion reduction. KBC’s experience is foaming and emulsion problems require some modification or upgrade to the high pressure separators to improve oil/water and vapor/liquid separation. As a final note, foaming in the amine contactor within the high pressure loop was excluded from this response, as this topic has been covered extensively in past NPRA proceedings.