During the 2015 American Fuel and Petrochemical Manufacturers Q&A and Technology Forum (Oct. 4-7, New Orleans), US domestic and international refiners dedicated time to discussing trends in hydroprocessing operations, with an extended focus on issues of safety, phosphorous poisoning, and meeting the US Environmental Protection Agency's (EPA) more stringent Tier 3 gasoline standards that take effect Jan. 1, 2017.
This annual meeting addresses real problems and issues refiners face in their plants and provides an opportunity for members to sort through potential solutions in a discussion with panelists and other attendees.
This is the first of three installments based on edited transcripts from the 2015 event. Part 2 in the series (OGJ, Sept. 5, 2016) will highlight discussion surrounding processes associated with crude-vacuum distillation and coking, while the final installment (OGJ, Oct. 3, 2016) will focus on fluid catalytic cracking (FCC).
The session included six panelists comprised of industry experts from refining companies and other technology specialists responding to selected questions and then engaging attendees in discussion of the relevant issues (see accompanying box).
The only disclaimer for panelists and attendees was that they discuss their own experiences, their own views, and the views of their companies. What has worked for them in their plants or refineries might not be applicable to every situation, but it can provide sound guidelines for what would work to address specific issues.
Safety
What are the likely causes for temperature excursion events in a hydrogen plant?
Epstein Hydrogen plant temperature excursions are possible in several of the catalyst vessels and are usually observed in association with the water-gas shift reaction. During normal operation, the high, medium, and low-temperature shift reactors display an exothermic reaction. Changes to feed quality can increase this exotherm and cause the operating temperature to exceed the maximum allowable working temperature. A critical alarm on a high-reactor discharge temperature is typically used to protect against this condition, with precursor alarms as needed.
A similar exotherm is observed during the reverse water-gas shift reaction over nickel catalyst in the methanator, which can increase under a change in feed quality. The exotherm generated by this reaction is significant enough that commonly occurring initiating events, such as the CO2-removal system's pump tripping, can cause the excursion. A safety integrity level (SIL)-rated system is typically used to isolate and depressurize the methanator if an excursion is detected.
I have personally observed an unexpected temperature excursion on startup of a combined feed pretreatment system with a catalyst containing nickel, molybdenum, cobalt, and zinc oxide used to combine qualities of hydrotreating and desulfurization beds into a single media.
During startup, the natural gas feed was introduced along with a very high hydrogen-recycle flow at normal operat ing temperature and pressure. The unit operating procedure called for starting the hydrogen-recycle flow at its full-flow rate followed by slowly ramping up the natural gas feed, which we did.
Under these conditions, it appeared that the hydrogen, CO2, and natural gas reacted exothermically through the reverse water-gas shift reaction, similar to the methanator. The catalyst also was very dry after prolonged nitrogen circulation. It appeared that the water generated by the reverse reaction was absorbed by the catalyst, causing it to heat up and initiate a reduction reaction that lead to the temperature excursion. Stopping the hydrogen flow discontinued all of this, and we were able return to a safe temperature level. According to the catalyst supplier, this was the first observation of such an issue with this material.
Wright We approached the question from a different angle. We recently experienced a temperature excursion caused by falling refractory in our transfer line. The transfer line is the pipe in between our reformer box and our first waste-heat boiler. The line consists of a metal shell with refractory on the inside. The top portion of the refractory fell, allowing the hot syngas access to the metal shell. The line quickly developed a fish-mouth rupture. Fortunately, no one was hurt, but there were cosmetic damages on the underside of the heater.
Other than that, in the past, we have just had out-of-the-pipe issues, such as leaky flanges around some feed heat-exchangers. We have had leaks around our pressure-swing adsorber valves, which caused fires. One time, we had a fire caused by corrosion and erosion from attemperator water used to control the inlet temperature to a shift converter.
Moreland At Valero, we have experienced high bed temperatures in hydrogen plants in the high-temperature shift reactor due to failure of the control system or control valve in the upstream process gas boiler resulting in high inlet temperatures. Additionally, the catalyst was very dry upon startup, with the heat of absorption resulting in temperatures of more than 1,200° F. In the methanator, as Paul mentioned, we have also had an issue with a failure in the upstream CO2-removal system. No catalyst damage, however, was observed in that experience.
Adkins I know this is not directly related to excursions in the reactor, but as far as temperature excursions in general, in one particular plant I experienced the failure of some of the safety systems related to a steam generator. The No. 1 process steam generator blew out because it ran out of water. So definitely keep in mind that you should be looking at level instrumentation and safety practices.
Phosphorus poisoning
Phosphorus-based chemicals are used to neutralize naphthenic acids. Drilling and completion fluids also can contain phosphorus, so it may be in crude oil. What are your best practices to protect active hydrotreating catalyst from phosphorus poisoning?
Moreland I am going to give a little background on phosphorous poisoning and then share one specific example we have seen in one of our refinery units. First of all, phosphorus is a strong catalyst poison. In our course materials, we say that 1 wt% phosphorous on catalyst will reduce activity by 50%, so it is a significant poison. There have been examples of up to 3 wt% phosphorous completely destroying the exotherm from a bed. The deactivation mechanism is pore-mouth plugging through surface deposition similar to nickel or vanadium. The common sources are crudes, drilling fluids, phosphorous-based corrosion inhibitors, and flow improvers. In particular, we have had more and more experience with phosphorous-based corrosion inhibitors, and the example I am going to discuss deals with that.
We have seen, in this example, the spent catalyst from the skim indicated up to 20 wt% phosphorous on top-bed grading. After seeing the exotherm decline during this cycle, when it came time to reload, we switched from the phosphorous-based naphthenic acid corrosion inhibitor to a sulfur-based inhibitor, which halted the decline in differential temperature (DT).
The unit in question contains two reactors in series with two beds each (Beds 1, 2, 3, and 4). Throughout the beginning of the first cycle, the DTs were stable, with the total rise in temperature distributed across the four beds.
When the phosphorous-based corrosion inhibitor was started, however, Bed 1's temperature rise started to decline. We tried to increase inlet temperatures to recover the DT, but it continued to decline. As you would expect from plug flow and metals deposition, Bed 1's declining DT eventually spread across Beds 2, 3, and 4.
We ended up changing the catalyst out in the cycle ahead of schedule. When the catalyst samples were pulled from the spent-catalyst load, Bed 1 catalyst samples exceeded 20 wt% phosphorous. While Bed 2 catalyst samples also exceeded our threshold of 1 wt% phosphorous, we did have lower levels of phosphorous in Beds 3 and 4, which is where the bulk of hydrodesulfurization catalyst was located.
At the beginning of the next cycle, the DT continued to decline across all four beds, which occurred during the time it took us to receive spent-catalyst sample results from the previous cycle. When we unload the catalyst and send it to our supplier, we do not get those results back for several months. This reinforces that having good technical service and timely return of spent-catalyst samples can really help us make a change.
As soon as we saw the high levels of phosphorous after receiving spent-catalyst samples back from the first cycle, however, we worked with our chemical supplier and the refinery switched to a sulfur-based corrosion inhibitor. Upon switching to that sulfur-based corrosion inhibitor, DT stabilized. While there was still some decline as a result of the rampant metals poisoning previously occurring in the unit, the rapid decline that we observed during the previous cycle was not repeated.
Our best practice for phosphorous poisoning, then, is to minimize the source of phosphorous poisoning. While we do still use phosphorous-based corrosion inhibitors for their effectiveness against naphthenic acid corrosion, there has to be an economic justification between their use for protecting vacuum gas oil (VGO) circuits and any deactivation that may occur in a VGO hydroprocessing unit.
There are demetallization catalysts that would be used for nickel and vanadium; they are also effective for trapping phosphorous. But still, a small amount of phosphorus goes a long way in terms of catalyst deactivation.
Temme That is a really powerful example. Phosphorous is definitely becoming more problematic, especially with treatment programs for crudes that have a high total acid number (TAN). Phosphorous will cause pore-mouth plugging leading to diffusion limitations. The mechanism is more pronounced with high-temperature, high-pressure operations. The pore-mouth plugging will have a greater impact on small pore-size, smaller median pore-distribution sized catalysts. We have seen that phosphorous deposition of 2 wt% can result in a 50% loss of catalyst activity. In terms of assessing and dealing with phosphorous trapping, as Andy indicated, getting good spent catalyst samples is essential for assessment. There are catalyst-handling companies that have core sample machines which work very well for getting a good axial core-bed sample.
Instead of having to try and vacuum off or pick samples out of flow bins, we want to make sure to get enough samples to define the extent of the problem. It does not hurt to get some third-party testing done, especially if they can do it in a quicker turnaround and get a baseline of the phosphorous and the catalyst substrate from the vendor to ensure that you have good accounting.
And yes, phosphates can be removed using guard catalysts. They are designed for iron, nickel, and vanadium pickup, especially catalysts with pore architecture that will minimize pore-mouth plugging. We feel that catalysts with minimal diffusion limitations will deal best with phosphorous trapping.
Pedersen I will add that it is important to identify not only the concentration of phosphorus contaminant but also the type of material with which this phosphorus is associated. The more acidic the phosphorus compound, the more lethal it is to your catalyst.
Johns I am curious about comments from the panel about protection from phosphorus in crude. Is there any monitoring program or other procedures you have instituted to be able to prevent that poisoning?
As an example, just so people are aware that this is a real problem, we had a diesel hydrotreater that completely went through a cycle in about 6 weeks due to phosphorus from crude poisoning. So what have you done?
McArthur I will say that we do not have a unit that has really had a significant phosphorus poisoning problem yet. So I do not think we have installed what I would call a top-of-the-line phosphorus monitoring program.
Moreland All I want to say, Jeff, is that we would have been caught by something like that, too. We do not typically monitor for phosphorus in diesel streams. We have seen some phosphorus in some kerosine (kero) streams at one of our plants and some phosphorus-based fouling in the crude tower in the kero circuit. So maybe it is a similar case, but it did not seem to affect operations. In that refinery, jet fuel is not hydrotreated, so we would not have seen it in the hydrotreater. We have monitored it in VGO, but in diesel, we do not look for phosphorus very often.
Appalla I have two specific questions. First, is there any industry-wide accepted limit of phosphorus in the gas oil stream going to the hydrotreater? Does anyone monitor this, and is there any permissible limit for the phosphorus levels?
Secondly, in the presentation, it was mentioned that the customer switched from phosphorus-based chemistry to sulfur-based chemistry. Phosphorus-based chemistry is known to provide a better protection against the attack of naphthenic acid by forming a more stable Fe-S-P scale. Sulfur-based inhibitors are not found to provide that level of protection. Also, as per experimental studies, the sulfur-based inhibitors need to be dosed in a higher proportion than the phosphorus-based to get the same level of protection. How did the customer ensure the same level of protection against corrosion from the naphthenic acid by switching to sulfur-based chemistry?
Moreland I can answer the second question. I do not know about the first one. In answer to the second part, we understood that switching to the sulfur-based inhibitor would be less effective for protecting against naphthenic acid corrosion. So the higher dosing went into the economic equation of how much high-TAN crude we process at this particular facility.
We have two different plants at which we are using the phosphorus-based corrosion inhibitor. One plant stayed with it and even dealt with the penalty on the hydrotreater downstream. The other plant, in the example I gave, switched away from the phosphorus-based to the sulfur-based inhibitor and then reduced its throughputs of high-TAN crudes. As far as acceptable levels, I do not really have a number.
McArthur The only time we review feedstock for phosphorus is if we have an opportunity feedstock come in, at which point we would look for elevated phosphorus levels. We would steer away from something that had any appreciable amount of phosphorus, but I do not have a number off the top of my head.
Prorok About fouling in the crude tower, at the Lima refinery, we experienced phosphorus fouling in a light gas oil cooler. We also experienced it in the "water white" kerosine cooler-a separate air cooler at the refinery-that basically plugged up with the same material and shortened the life of the diesel hydrotreater catalyst.
Moreland We saw it in deposits in the crude tower. I do not know if we knew it ahead of time, because when we analyzed those deposits, that is when we saw phosphorus in those deposits. Is that a similar experience?
Prorok After the pressure dropped in the kerosine section of the crude tower, we retrieved a sample of the solids that were caught in a strainer on a pump, had them analyzed, and found out that there was phosphorus in that circuit. Then, we saw phosphorus in the heat exchangers downstream of that, on the way to the distillate hydrotreater (DHT).
Price There is a question on this in the Crude Q&A that covers phosphorous deposits in the crude unit equipment that may provide additional insight.
Torrisi The panel talked about solids in the crude tower, and the focus of the question was activity on the catalyst. Can anyone on the panel comment about any impacts on pressure drop related to phosphorus?
McArthur I do not have any experience with it, but my understanding is that there will be a pressure drop problem with phosphorus fouling on a catalyst, especially if you have a very low void space where it is collecting. But again, we have not had that problem yet.
Torrisi The experience I have seen has been more with regard to accelerated fouling in the feed preheat exchangers than in the catalyst bed itself.
Moreland Sal, you are probably as familiar with the unit as I am, so all I will add is that when you lose the activity this fast from phosphorus, the time on stream is not sufficient for development of a significant pressure-drop issue. We have seen pressure drop start to increase, however. We ended up doing reactor skims before it really became a run limiter.
Torrisi Okay, thanks. I was wondering about that because what happens is you use a corrosion inhibitor to neutralize the acids, and if you are not consistently neutralizing, organic iron can form as a result of the reaction of the acids with the equipment.
We have observed organic iron being converted back to solid iron sulfide by reaction with H2S in the reactor at the same time we observed the phosphorus being deposited in the catalyst bed. That is a secondary indicator that you may have a phosphorus issue because you are just starting to see some iron sulfide formation due to conversion of iron naphthenates, which can go hand-in-hand with phosphorus.
Tier 3
How is your company planning to meet Tier 3 gasoline regulations?
Wright Currently, our gasoline runs at 12 ppm sulfur. There are two sour components: our light straight-run (LSR) and butane. Some approaches we are contemplating for compliance are hydrotreating the LSR, reducing the sulfur via dilution, or restarting an out-of-service Merox unit.
Epstein Flint Hills Resources includes two refineries. Our Pine Bend refinery will increase the amount of hydrotreating in gas oil hydrotreaters to remove sulfur from fuel products to meet Tier 3 regulations. Our Corpus Christi refinery is reconfiguring its ultralow-sulfur gasoline unit in order to meet these requirements.
McArthur Phillips 66's compliance strategy varies by refinery. Our West Coast sites are generally already at a Tier 3-performance level. We have some unit revamps to allow for higher severity operation of our FCC-feed hydrotreaters. We have some new Merox units we are installing, and where necessary, we are adding some additional hydrotreating capability.
Moreland For the 14 Valero refineries, three are not affected by Tier 3, either in Europe or California. Prior to Tier 3, Valero met gasoline sulfur requirements with just FCC-pretreat at three of our refineries that process light sweet crudes.
For two of those three refineries, we are going to build grassroot gasoline desulfurization units, while at the third, we have shut down the FCC. So for compliance with Tier 3, all Valero refineries will have FCC gasoline hydrotreating or post-treatment units. I also would like to add that for Tier 3 gasoline, a lot of our front-end studies have had to do more with how we deal with crude-unit naphtha than with FCC-gasoline naphtha. In many of our cases, the crude-unit naphtha is not hydrotreated to meet the 30 ppm pool but must now be hydrotreated to meet the 10 ppm pool. Tier 3 will require refiners to hydrotreat a lot of straight-run naphtha.
Varraveto My question regarding Tier 3 compliance has to do with the strategy or use of credits and how that fits in for any of the operating companies represented on the panel or in the audience.
Moreland Since the EPA has allowed us to use 5 years' worth of credits, Valero has sufficient credits built up to maybe delay startup of those two grassroots units I referred to earlier.
Larson Considering not only Tier 3 but also the emissions requirements, I am curious why more FCC-feed hydrotreating will not be the preferred option. Feed pretreating improves stack emissions compliance as well as fuel side issues.
It seems like that is the direction the industry is moving: to zero flue-gas sulfur that you get two for one (i.e., lower emissions and increased hydrogen content of feed, which improves conversion flexibility and selectivity). And if you are building new FCC-pretreat capacity, it will justify expense on a per-barrel basis vs. the lower pressure post-treatment units.
Wildenberg We found that managing hydrogen recycle and hydrogen purity can keep our deactivation rates low. Even though we are more severely hydrotreating, it seems like our crude people keep finding harder and harder things for us to hydrotreat. So at the Pine Bend refinery, we have been enjoying very low deactivation rates. But, like I said, it requires a lot of management, catalyst strategy, and careful management of higher hydrogen purity-higher hydrogen recycle.
I also have been asked about run length and product sulfur. One example is our heavy coker, heavy vacuum gas oil (HVGO) hydrotreater, that has minimal catalyst deactivation. This unit shut down after 3 years at the scheduled turnaround, but it could have gone longer. It has been averaging 0.05-0.07 wt% product sulfur with 3.5 wt% sulfur feedstock, so it is doing well.
Moreland Do you still get all the way to 10 ppm on FCC gasoline with feed from that cat feed hydrotreater?
Wildenberg We get to 30 ppm or less on the FCC gasoline from that feedstock. We have two gas oil hydrotreaters that hydrotreat 100% of the gas oil the Pine Bend refinery runs. The lighter feedstock goes to the 900-psig unit, and the heavier feedstock goes to the 600-psig unit. So we managed these units and their catalyst strategies very diligently.
Moreland In our refinery in Wilmington, Calif., Valero is running similar coker gas oil and HVGO to produce FCC feed with 500-700 ppm sulfur content. We typically only get a 2-year cycle.
Wildenberg We went back and looked at a lot of the deactivation times we had been doing and discovered that we had done a lot of it to ourselves. We were either cutting hydrogen too far or raising temperature too much. So we really steadied down our temperature moves. We do not target the product sulfur and move temperatures around. We allow the sulfur to move around some to keep the temperature very steady, and then we are able to not deactivate. It was definitely helpful.
The panel
Paul Epstein, senior process engineer lead, Flint Hills Resources LP
Scott McArthur, senior hydroprocessing engineer, Phillips 66
Dr. Andrew Moreland, hydroprocessing and hydrocracking technology advisor, Valero Energy Corp.
Mike Pedersen, senior technical service specialist, UOP LLC
Paul Temme, hydroprocessing technical supervisor, Albemarle Corp.
Samuel Wright, senior process engineer, Hunt Refining Co.
The respondents
Mike Adkins, KP Engineering LP
Jeff Johns, Chevron USA Inc.
Nagashyam Appalla, Reliance Industries Ltd.
James Prorok, Husky Energy Inc.
Maureen Price, Fluor Corp.
Sal Torrisi Jr., Criterion Catalysts & Technologies LP
Dominic Varraveto, Burns & McDonnell
Mel Larson, KBC Advanced Technologies Inc.
Wendy Wildenberg, Flint Hills Resources LP
