HP_2010_06

JUNE 2010

HPIMPACT

SPECIALREPORT

TECHNOLOGY

Steam-cracking furnace modules

PROCESS AND PLANT OPTIMIZATION

Avoid engineering problems

EPA weighs decision on E15

New strategies focus on energy, reliability profits and environment

Improve emergency depressurization procedures

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JUNE 2010 • VOL. 89 NO. 6 www.HydrocarbonProcessing.com

SPECIAL REPORT: PROCESS/PLANT OPTIMIZATION

35

How will carbon emissions regulations revise energy conservation economics?

43

Low-cost advanced process control project captures energy savings in utilities area

Including the cost of carbon in refinery project economics has the potential to convert previously marginal energy projects into more attractive options I. M. Glasgow, S. Polcar, E. Davis, T. Nguyen, J. Price, C. Stuecheli and R. E. Palmer

This low-cost project resulted in over $300,000 per year in benefits E. Chang and M. Viducic

49

Modify your vacuum-tower transfer line to increase benefits

55

Advanced process control: quick and easy energy savings

58

Consider low-voltage AC drives for hazardous areas

61

Designing the vacuum unit is a challenge. Revamping the transfer line can make this unit more profitable.

R. Yahyaabadi Here are some of the energy savings benefits routinely produced by implementing advanced process control P. Kesseler Such drives used with electric motors can offer significant process improvements and energy savings for equipment used in explosive zones J. Riikonen

Trading silicon for carbon: how to reduce energy usage through automation The average plant can conservatively achieve 15% energy savings through this technology D. C. White

SAFETY/LOSS PREVENTION

71

Cover Round-the-clock operation of the integrated transport network is one of Gassco’s main jobs as operator. These activities are run from the control center at the company’s head office, which lies at Bygnes in Karmøy local authority north of Stavanger. The main duties are to manage the gas flow through constant monitoring, regulate quality and ensure that the gas blend is correct. Since the various fields deliver different gas grades, these must be processed or mixed to achieve the desired quality before delivery to the buyer. Such blending allows us to deliver gas with the agreed composition virtually all the time.

HPIMPACT 19 Large steam-cracking furnace modules arrive in Singapore 19 EPA urged to delay judgment on E15 20 Energy future must include mix of renewables

Performing correct initial blowdown calculations These guidelines can help determine peak flowrate A. Vyas

PLANT DESIGN AND ENGINEERING

75

How to avoid interface engineering problems Here are several ways to accurately exchange information S. Moulik and K. Saito

HEAT TRANSFER/RELIABILITY

77

Troubleshooting waste-heat boiler poor heat transfer Analytical results showed that potash leached from the catalyst, refractory powder, corrosion products and, to a lesser extent, catalyst fines resulted in tube fouling G. Yeh, I. Al-Babtain, S. Al-Zahrani and N. Al-Ghanemi

COLUMNS 9 HPIN EUROPE Technology goliaths vie with startups to crack biojet processes

DEPARTMENTS

11 HPINTEGRATION STRATEGIES Energy management is a work-in-progress

7 HPIN BRIEF • 13 HPIN ASSOCIATIONS • 25 HPINNOVATIONS 29 HPIN CONSTRUCTION • 32 HPI CONSTRUCTION BOXSCORE UPDATE 82 HPI MARKETPLACE • 85 ADVERTISER INDEX

15 HPIN CONTROL Total- or partial-draw configuration

HP ONLINE EXCLUSIVES

17 HPI VIEWPOINT When it comes to information technology, the sky’s the limit

Purging and inerting large-volume tankage and equipment—jet mixing concept—Part 2 Here are the advantages and disadvantages of the various methods M. Gollin

Implementing a suitable safety instrumented system—Part 2 The analysis is the most important step for engineering and designing a suitable system R. Modi

Furnace tubes—enhancing safety and efficiency with infrared thermography Thermal imaging offers advantages over thermo couples M. Cronholm

86 HPIN WATER MANAGEMENT What’s happening with the Legionella Standard?—Part 2

Italian design A masterpiece

www.HydrocarbonProcessing.com Houston Office: 2 Greenway Plaza, Suite 1020, Houston, Texas, 77046 USA Mailing Address: P. O. Box 2608, Houston, Texas 77252-2608, USA Phone: +1 (713) 529-4301, Fax: +1 (713) 520-4433 E-mail: editorial@HydrocarbonProcessing.com www.HydrocarbonProcessing.com Publisher Bill Wageneck bill.wageneck@gulfpub.com EDITORIAL Editor Les A. Kane Senior Process Editor Stephany Romanow Process Editor Tricia Crossey Reliability/Equipment Editor Heinz P. Bloch News Editor Billy Thinnes European Editor Tim Lloyd Wright Contributing Editor Loraine A. Huchler Contributing Editor William M. Goble Contributing Editor Y. Zak Friedman Contributing Editor ARC Advisory Group (various) MAGAZINE PRODUCTION Director—Editorial Production Sheryl Stone Manager—Editorial Production Angela Bathe Artist/Illustrator David Weeks Manager—Advertising Production Cheryl Willis ADVERTISING SALES See Sales Offices page 84.

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HPIN BRIEF BILLY THINNES, NEWS EDITOR

BT@HydrocarbonProcessing.com

The climate change bill recently introduced in the US Senate should be rejected, according to a statement from the National Petrochemical and Refiners Association (NPRA). “The draconian carbon reduction targets and timetables in this bill would trigger destructive change in America’s economic climate,” the NPRA said. “This would add billions of dollars in energy costs for American families and businesses, destroy the jobs of millions of American workers, and make our nation more dependent on foreign energy sources.”

Meanwhile, illustrating a difference of opinion within the industry, Shell Oil Co. issued a statement in support of the bill. “This legislation ensures America’s global competitiveness and recognizes the role clean natural gas can play in growing the economy and protecting the environment,” Shell’s statement said. “[This is] a fresh approach to transportation emissions that encourages the development of home-grown energy, provides transparency for consumers and enables American refiners to compete against imports of diesel and gasoline.”

The sustainable oil price range for the next two years is $70–90 per barrel, according to ESAI’s two-year outlook for the global oil markets. The marginal cost of crude oil, such as Canadian oil sands, encourages a floor around $70, and spare refining and oil production capacity encourages a ceiling around $90. More importantly, the report points out, non-OPEC production should be stronger in the coming years than recent history would suggest. Along with impressive growth in OPEC NGLs and alternative fuels, the call on OPEC crude will only grow slightly.

The global engineering and construction (E&C) sector is off to a fast start in 2010, according to a new PricewaterhouseCoopers report. Deal value in the sector totaled $21.2 billion in the first quarter of 2010, compared with $3.8 billion and $23.7 billion in the first and fourth quarters of 2009, respectively. There was also a strong rebound in mega-deals (transactions of at least $1 billion) during the first quarter of 2010, with four transactions announced. Conversely, no mega-deals were announced in the first quarter of last year and there were only seven in all of 2009.

Targets located in North America, Asia-Pacific and Europe were the primary drivers of deal activity in 2010’s first quarter. Transactions involving both US targets and buyers were the key drivers of deal activity, particularly in terms of deal value. Of the 29 transactions announced during the period, nine (31%) involved a US entity. Of the $21.2 billion in deal value announced during the period, 61% was attributable to US-affiliated activity.

API presented environmental and safety awards to several US pipeline companies at its 61st Annual Pipeline Conference in New Orleans, Louisiana.

CITGO Pipeline Co., Genesis Energy, Inc. and Western Refining Pipeline Co. received API’s Occupational Safety Performance Award for the small operator category. ExxonMobil Pipeline Co. received the award for the large operator category.

Portland Pipe Line Corp. and Alyeska Pipeline Service Co. took API’s Environmental Performance Award in the small and large operator categories, respectively. The Occupational Safety Award is presented to companies with the lowest OSHA recordable incidents for employees and contractors. The Environmental Performance Award is presented to companies with a track record of stellar environmental performance.

Led by Brazil, Latin America is expected to reach 46 GW of total installed wind capacity by 2025, with a 12.6% compound annual growth rate of yearly installations, according to a new market study from IHS. Brazil will lead the region with 31.6 GW installed by 2025, trailed by Mexico with 6.6 GW. Chile will also add significant wind power, boosted by the country’s renewable portfolio standard, according to the study. HP

■ US demand for refinery chemicals US demand for refinery chemicals is forecast to exceed $7.5 billion in 2014, according to a new industry study from The Freedonia Group. Advances in chemical demand are expected to slow significantly from the rapid pace achieved during the 2004–2009 period.

Decelerating growth will be based on an outright decline in refined products output, combined with the completion of the phase-in of a 30-ppm sulfur limit in gasoline, enacted in 2004. However, overall demand for refinery chemicals will benefit from an expected acceleration in economic growth, as the economy eventually recovers from the recession that began in late 2007. Increasingly stringent environmental regulations will further promote demand, as refiners subject their products to higher levels of chemical treatment in order to remove impurities. Market growth will result primarily from above-average gains in the large merchant hydrogen segment, due to rising use by refiners seeking to supplement their captive hydrogen production.

Merchant hydrogen will remain the largest and fastest growing product in the US refinery chemical market.

Advances will be driven by tightening sulfur standards for diesel fuels.

Such environmental regulations promote the use of hydrotreating as a means of removing sulfur and other contaminants. As of 2010, diesel fuel must meet a 15 ppm sulfur limit, and going forward, significant sulfur reductions are also expected for heating oil. Hydrocracking represents another growth application for merchant hydrogen, as US refineries continue to expand their hydrocracking capacity in efforts to boost gasoline and diesel fuel yields.

Metal catalysts will maintain their position as the largest and fastest growing segment of the refinery catalyst market. Gains will be based on rising use in hydrotreating applications due to efforts to reduce sulfur content in refined products. Zeolites represent another leading type of catalyst.

Primarily employed in catalytic cracking applications, the relative maturity of this technology will serve to limit gains for zeolite catalysts. HP HYDROCARBON PROCESSING JUNE 2010

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HPIN EUROPE TIM LLOYD WRIGHT, EUROPEAN EDITOR tim.wright@gulfpub.com

Technology goliaths vie with startups to crack biojet processes This month, I had a chance to verify my armchair thesis that all liquid biofuels will end up in the air—diverted to the passenger aviation industry and the world’s air forces. At a fuel forum attended by members of the international aviation industry in Los Angeles (LA), California, there were presentations and a small “marketplace” for biojet fuel technologists who were showing their wares. Biojet actors will get a higher-profile showcase at the UK Farnborough International Air Show later this year.

R&D for biojet fuel technologies. A massive R&D effort

is already preoccupying development companies such as Honeywell’s UOP, Rentech and, through a recent $600 million acquisition, ExxonMobil. And this year, European oil companies will have to decide when and how they are going to address requirements to comply with renewables and greenhouse gas reduction directives. These are likely to mean a substantial and growing price advantage to fuels that are blended with renewable elements.

Up in the air. This is just the start. It seems likely that the aviation market, in time, will suck up all of a fairly limited, though admittedly renewable, liquid biofuel resources. There seems to be no other alternative. Apart from the obvious alternative of collectivizing personal transport and simply not shipping so much stuff by localizing and deglobalizing production, there are fuel alternatives for ground transportation.

In February 2010, I discussed that ships are now being operated with hybrid compressed natural-gas and fuel-cell drives.

Perhaps, dimethyl ether, from gas or biomethane, can compliment diesel in the trucking fleet. For cars, there is biomethane, conventional electric, lithium electric and, apparently, compressed air. But none of these fuel sources work well.

For their part, the airlines have been busy recently signing memos with producers and running test flights to help these nascent air fuels toward certification. The airline industry is still spooked by events in 2008, when jet-fuel prices climbed to almost double today’s price. Thus, they see a practical solution in biojet.

Biofuel efforts also help demonstrate opportunities to policy makers, who are eager to embroil these “green” fuels into Europe’s ambitious renewables and emissions trading programs.

But leading the civil aviation sector is a well-funded R&D powerhouse—the US military. Stemming from acute interest in biofuels, a small Swedish R&D outfit has been swept into the fast-track by a DARPA financial award.

Angelica Hull is the managing director of Swedish Biofuels AB. This company is using its technology to produce diesel blends for standard engines in its home country. While in LA, Swedish Biofuels AB was showing its isoparaffinic jet fuel process. Appropriately for Sweden, its technology uses timber as a feedstock. A proprietary process breaks down ligno-cellulose into higher alcohols and then

into hydrocarbons that can be blended at up to 100% as road and jet fuels. Ms. Hull is skeptical about a plan at the Stockholm Airport to use a Fischer Tropsch-based biogasification plant to provide fuel for the airport. “Using our process, you can produce just as much fuel for a fraction of the investment cost,” she said.

Following development under DARPA funding, UOP’s Green Jet Fuel looks like it will be the first renewable to be fully certified for aviation applications. “We were able to use a lot of what we learned in developing our Ecofining process with ENI for the production of Green Diesel,” Graham Ellis of UOP said. UOP had coined the term Green Diesel to differentiate the product of its process from biodiesel, which had come to represent a specific molecule.

“Then DARPA came to us and asked us to develop a Green Jet Fuel,” he explained. “In the process for our Green Jet Fuel, we not only have to isomerize, but also to crack,” Graham explains.

“We’re now making on-spec bio-synthetic paraffinic kerosene (SPK) with a freeze point of –47 that involves some fairly severe processing,” he said.

Path forward. UOP, Swedish Biofuels and other technology

companies are well along the path to certifying their fuels for use in civil and military applications. “It took 15 years for Fischer Tropsch jet fuel, so we’ll just have to see how we get on,” said Ms. Hull, whose company is just about to start trials with the Swedish and US air forces.

Browsing the UOP stand at the forum was Michael Lakeman, Boeing’s Seattle-based regional director for biofuel strategy. “The technology is ready to go,” Lakeman said. “It’s securing a costeffective, demonstrably sustainable supply of feedstock—and at scale—that is at issue.”

But what I really wanted to know from Boeing was if it has a game changer in the wings. For many years, the car and oil industries reaffirmed that alternatives must be “drop-in” liquid fuels requiring no changes to fuel stations or drive trains. Suddenly, albeit with the lithium supply chain in question, automakers globally are committing to developing electric vehicles and hybrids.

“We have people at Boeing whose job is to look up to five decades into the future and they themselves are telling us that there really is no alternative for air travel than energy-dense liquid fuels,” Lakeman said. “We just don’t have any alternatives,” he added, “ground transportation does.” Lakeman added that companies like his believe the air industry should have preferential access to biofuels. “For ground transport, there are other options, but aviation has nowhere else to go,” he said. HP The author is HP’s European Editor. He has been active as a reporter and conference chair in the European downstream industry since 1997, before which he was a feature writer and reporter for the UK broadsheet press and BBC radio. Mr. Wright lives in Sweden and is the founder of a local climate and sustainability initiative.

HYDROCARBON PROCESSING JUNE 2010

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HPINTEGRATION STRATEGIES ALLAN AVERY, CONTRIBUTING EDITOR allan.avery@arcnet.com

Energy management is a work-in-progress Energy accounts for a significant portion of overall manufacturing costs. With today’s volatile energy prices, HPI and other manufacturers have to keep a close eye on where they get their energy and how much energy they use. An ongoing ARC survey indicates that many owner-operators already have strong energymanagement initiatives in place and employ automation and control technologies to monitor and optimize plant energy use.

However, some respondents indicated that they have limited visibility into energy-management metrics, and many do not measure and control energy usage in real time, leading to missed opportunities to reduce energy costs.

Energy is often the largest component of an HPI or other manufacturer’s cost structure. According to the EIA Annual Energy Outlook for 2010, energy prices will increase substantially over the next couple of decades, and efficiency alone will not be enough to stem manufacturers’ rising energy costs. Plant operators will have to do more to realize additional savings, including realigning their energy-management practices to optimize energy use and costs in a climate of expensive energy.

Since September 2009, 150 individuals have responded to ARC’s ongoing energy management survey. Respondents represent a broad spectrum of manufacturers and equipment suppliers from the process (34%), discrete (43%) and hybrid (23%) manufacturing industry sectors. Corporate management, engineering and marketing functions are well represented. Most respondents are located in North America and Europe.

Energy-management initiatives. Most survey respondents have energy-management initiatives in place and their programs are well integrated with both business and production processes.

Nearly 60% of respondents have an “energy czar” or leader, more than a third who are at the top levels of the organization.

More than half of the respondents have an active energy-management team to handle energy-efficiency efforts and projects. For Compressors

58.8%

Power generation

39.7%

Steam generation

36.8%

Fired heater

25%

Boilers

29.4%

Others 0%

FIG. 1

22.1% 20%

40%

Main plant energy challenges.

60%

80%

nearly half of those, the automation group plays a role in energymanagement efforts. At these companies, energy management factors into day-to-day plant operations.

Energy and production costs. Among the survey respon-

dents, roughly one-third reported that energy makes up 20% or more of their overall production costs, and 40% of respondents indicated that energy accounts for between 10% and 20% of their production costs.

While energy makes up a significant portion of production costs, less than a third of the respondents actually monitor energy costs in real time, while more than half look at energy costs periodically for cost allocation and accounting purposes. Fifteen percent of respondents do not monitor production-related energy costs.

Energy pain points around the plant include the usual suspects, such as boilers, fired heaters and steam plants. Respondents also reported that compressed air was a primary energy challenge.

Nearly 60% indicated that compressors hog plant energy, while in-plant power and steam generation also pose significant challenges (Fig. 1).

Energy KPIs. A sizeable portion of respondents use overall

plant energy consumption as a KPI. However, almost half opt for more granular metrics that tie energy use to individual products or product revenues. A small, but significant, group measures success through reduced CO2 emissions. With today’s increased emphasis on sustainability and environment, it’s likely that new or additional carbon tax or cap-and-trade legislation will be enacted in the near future. This will tag on additional production costs, and manufacturers that already monitor carbon emissions will have a jumpstart on managing carbon costs.

Technology and energy management. Many respon-

dents apply automation and control technologies to help reduce energy consumption in their plants, particularly intelligent pumps, simulation, decision support, combustion analyzers and intelligent field devices. However, penetration in plants is relatively shallow; most reported using these technologies in between 5% and 15% of their process units. This low usage rate could indicate that respondents are still in the pilot project stage, but automation technologies could be more broadly applied for energy management, and that users are leaving potential energy savings on the table.

Larry O’Brien is part of the automation consulting readers team at ARC covering the ARC encourages Hydrocarbon Processing to participate process industries, and an HP contributing editor. He is responsible for tracking the in our ongoing energy management survey at: http://www.arcweb. market for process automation systems (PASs) and has authored the PAS market studcom/research/surveys.

HP has also authored many other market research, ies for ARC since 1998. Mr. O’Brien strategy and custom research reports on topics including process fieldbus, collaborative partnerships, total market trendsGroup, and others.

He has been with since The author , anautomation analyst at ARC Advisory recently completed anARC extensive January started his career withpractices. market research in the field instrumentation end-user1993, studyand on energy-management He graduated from the University markets. of Rhode Island.

HYDROCARBON PROCESSING JUNE 2010

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editorial@HydrocarbonProcessing.com

The good, the bad and the ugly The good. Offshore Northern Seas (ONS) offers a biennial press trip to 20 editors across the world. Pramod Kulkarni, World Oil’s new editor, and I were invited to attend. The meeting started in Stavanger with talks ranging from the global energy outlook, the Norwegian Continental Shelf (NCS), deepwater drilling and how Statoil is a reliable gas supplier to Europe. The following day, we took a ferry to Hywind, Statoil’s first full-scale floating wind turbine. The fact that Statoil modeled wind and ocean current waves that could affect the structure was unique. Then we were off to Gassco’s headquarters.

Brian Bjordal, CEO and president of Gassco, gave a presentation on Norway’s position as a reliable gas exporter. He focused on the global energy challenge and how Gassco’s gas processing plant is a safe and efficient gas transporter. We learned that Norway, from a gas production vs. consumption ratio, produces 100% but only consumes 5% for 2008. In 2007, Norway relied on electricity from renewable sources at 100%. Norway delivered 96.6 billion cubic meters of natural gas to the European Union in 2009.

Mr. Bjordal stated that “solving the world’s energy problem is very important and isn’t an easy task, but the good news is, that natural gas will be part of the solution.”

This will put pressure on coal use, since oil and gas is a substitute for coal. That there are significant reserves of conventional gas along with the newly discovered unconventional shale gas are positive aspects. What is unconventional gas all about? Given the assumption that gas is a component in the energy solution, gas—from a volume, depletion and distribution perspective in the world—is seeing positive news.

Mr. Bjordal commented that windmills are an interesting component to the energy solution; however, they don’t work when the wind doesn’t blow. “How do you have a sophisticated energy system that supplies a security of supply in extreme situations,” retorted Mr. Bjordal. “There has to be a way to build an infrastructure

that handles all these components.” One component he’s really concerned with is CO2 emissions, and Mr. Sigve Apeland gave an in-depth discussion on carbon capture and storage (CCS) technology at Kårstø and Mongstad.

The CCS technology Gassco is working on is amine and chilled ammonia through the European CO2 Technology Center in Mongstad and a plant is under construction. Unfortunately, this technology has not proven to be cost-effective. A research center was established for CCS that involves an eight-year research program. The program looks at transportation and storage solutions, and handling CO2. Several locations have been evaluated for subsea geological storage. Pipeline infrastructure is being considered in place in the NCS. After the presentations at Gassco headquarters, we were bused to Kårstø (Gassco’s gas processing facility) for more discussions and a plant tour. Ulf Rosenberg, PR manager for GDF Suez, talked about how the company is a new operator on the NCS and is building for the future. GDF Suez is a world leader in liquefied natural gas (LNG)—the No. 1 importer and buyer in Europe and the No. 2 operator of LNG terminals. GDF Suez owns the largest number of reserves (excluding the majors). Snøvit was the first LNG project in Europe and it can produce 4.3 million tons of LNG a year at full capacity.

Mr. Rosenberg gave us a history lesson on NCS production. Around 1999–2000, the Norwegian authorities were concerned

Gassco’s gas processing plant at Kårstø.

with many of the major mergers and started losing competition on the NCS. However, the majors started pulling out once production fell since that area wasn’t suitable for the discovery of new “elephant fields” and they put their resource/exploration money in other areas. At that time, only a third of the NCS reserves were produced, a third were planned to be produced and a third were yet to be discovered. How were the Norwegian authorities going to get this last third of reserves out? To increase the competition, the Norwegian Petroleum Directorate (NPD) CEO held road shows to invite independent oil companies and utilities to urge them to come to Norway and look into the resources still available.

The NPD CEO was successful and was able to get 50 new oil companies to start production in the last decade.

The bad. Unfortunately, our meeting with Norway’s Minister of Petroleum was canceled due to Iceland’s volcano eruption.

Also canceled was our scheduled helicopter ride from Haugesund to Kårstø.

The ugly. A second volcanic plume post-

poned our flight to Houston by six days.

Thanks ONS for your dedication to getting us back home. HP

Statoils’s Hywind project.

HYDROCARBON PROCESSING JUNE 2010

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Select 91 at www.HydrocarbonProcessing.com/RS

HPIN CONTROL Y. ZAK FRIEDMAN, CONTRIBUTING EDITOR Zak@petrocontrol.com

Total- or partial-draw configuration Ever wonder why some main fractionators are designed with partial-draw trays, whereas others are total draws? A typical partial-draw with pumparound configuration is shown in Fig. 1. Liquid is taken from the draw tray into a side stripper, and stripped, typically by steam, to become a side product. Pumparound is a device for removing heat from the column. It circulates liquid from the draw tray, via heat exchange, back to the column several trays above the draw. The now subcooled liquid exchanges heat by condensing vapor on these three or four trays, and thus, the pumparound section trays are used solely for heat exchange, having no separation effect.

The total internal reflux coming down onto this tray is higher than the pumparound plus product draw, and excess internal reflux flows over the weir down to the tray below, hence, the name: partial-draw tray. The excess internal reflux must be significant, or else the section below the draw tray becomes ineffective.

Pumparound heat removal is typically controlled by changing pumparound flow, and if the cooling circuit is not against a heat-availability limit that is an effective control mechanism.

Side-product flowrate is set by the operator, or it could be an APC manipulated variable, whereas the draw is on stripper level control.

The challenges APC engineers face are in determining how much heat is to be removed, how much side product to draw, what to do when the cooling circuit is indeed against a heat removal pinch limit and how to estimate internal reflux below the draws. Estimating the internal reflux from heat balance is imprecise because it involves subtracting two big numbers: condensation due to pumparound cooling minus product draws, and the precision gets worse going down the column.

Can this uncertainty be improved? Fig. 2 shows a totaldraw configuration, where all of the liquid on the draw tray is taken out to a draw pot. The pumparound pump in this configuration handles not only pumparound material but also pumpdown of internal reflux. Excess material not pumped from the pot is taken into the stripper and becomes the side product. The main improvement here is that internal reflux is being measured and there is no longer a need to estimate it. Another difference is the indirect manipulation of side draw. To increase side draw, the operator or APC must decrease pumpdown.

And there are other improvements. Holding a steady internal reflux profile on the column stabilizes the temperature profile and, in turn, product cutpoints. Further, total-draw tray configuration usually incorporates a temperature point just below the draw tray, which helps with the cutpoint inference. There is no reason why part-draw configurations would not incorporate a similar temperature point, but for some odd reason that is not traditionally common.

I come now to the original question. Have you ever

wondered why draw trays on some units are partial whereas on

FC

Side stripper

FC

Pumparound

LC FC

FIG. 1

Partial-draw configuration.

Pumparound FC

FC

Pot TI

Stripper

LC

FC

Pumpdown LC FC

FIG. 2

Total-draw configuration.

other units they are total? Crude-unit atmospheric fractionators are mostly part draw, while vacuum fractionators are by and large total draw. FCC and coker fractionators often incorporate mixed designs, with some draw trays being part draw while others are total draws. Is there any rhyme or reason? If total draw designs are easier to control why aren’t they implemented?

What about mixed designs? I would say that even if most draws are partial, configuring a lower-draw as a total-draw tray does offer advantages. Some crude units are configured with the lowest side draw: HGO as a total-draw tray and that helps stabilize the overflash. On the other hand, I cannot see how an upper total-draw tray, while the lower draws are from part-draw trays might be useful. HP The author is a principal consultant in advanced process control and online optimization with Petrocontrol. He specializes in the use of first-principles models for inferential process control and has developed a number of distillation and reactor models. Dr. Friedman’s experience spans over 30 years in the hydrocarbon industry, working with Exxon Research and Engineering, KBC Advanced Technology and since 1992 with Petrocontrol. He holds a BS degree from the Israel Institute of Technology (Technion) and a PhD degree from Purdue University.

HYDROCARBON PROCESSING JUNE 2010

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Call +1-713-428-5100 Visit www.merichem.com/ptd Email ptdsales@merichem.com Select 79 at www.HydrocarbonProcessing.com/RS

HPI VIEWPOINT When it comes to information technology, the sky’s the limit Weak demand and plummeting margins? The answer for greater productivity could reside in your current IT system—and the sky.

Mike Bannon leads the US Energy and Process Industry solutions organization at Microsoft. His team collaborates with Microsoft’s customers and partners to develop technology visions, architectures and applications that improve business performance. His experience spans 30 years, serving the oil and gas, electric utilities and process industries with industry leaders including Schlumberger, Honeywell and SAP.

Talk to me. In this example, we could imagine that at the start, the refinery workers were like persons together in a row boat.

They all knew their destination but weren’t rowing together.

Business insight and collaboration got everyone rowing together to reach their goal. What they didn’t know was that they had materials in the boat to make a sail that would have helped them go faster. That’s where the next level of collaboration comes in and where we can take a lesson from what’s happening on the Read any headline and it becomes clear: Cost cutting has Internet today. become the new normal in the HPI sector of the economy. RecesPopular social sites such as Facebook and Wikipedia are spreadsion-battered refiners and petrochemical producers have faced ing knowledge rapidly about many things, much of which may be weak demand and plummeting margins; some have been forced to inconsequential. But when a model like this can be used to make close plants and lay off workers. Many industry observers expect social connections and spread knowledge this economic pressure on the HPI to consecurely around a company, the power tinue for at least two years. The stakes have ■ New capabilities that you to tap the knowledge of experts within never been higher for industry to innovate and across the business partner ecosysin order to improve productivity. can use to transform your tem can help identify opportunities like So, your CFO asks, what have you adding a sail to the rowboat. That same done today to lower our breakeven cost? business today are out there. refiner created a global community with For many, the answer could lie within their collaboration software that solves simple and difficult issues by current information technology (IT) systems and new ones in the linking expertise to problems regardless of location. In only a short sky—more about that later. time, millions of dollars in avoided costs were documented.

One version of the truth. For a start, we can all agree on what to measure, how to measure it, how to set a target for it and Look up. Yes, look up. There are clouds overhead—computing how to communicate it. That’s where leaders of a multinational clouds. Imagine a processing plant that can instantly scale to cope refiner started on their quest to improve productivity. Having with from zero to extremely large demand and back again, on a derived and published a comprehensive set of metrics and opermoment’s notice. That plant, built by someone else, is available ating targets, they still found that productivity wasn’t zooming to you 24/7 and only charges based on the capacity you consume. ahead. So while one version of the truth is essential, it’s only the Pretty good deal, eh? Cloud computing, in a utility sense, works starting point. Ease of access to data and presentation in a way like that. You can get computing power on demand at a very that relates it to the user’s job function or role, as well as collaboralow cost without capital expense in a private cloud or the public tion among the users, turns out to be important, too. cloud. But a computer without data is like a plant without feedLow-cost and familiar software available on the desktop today stock. You just can’t get a lot done. So, the cloud has to store your can automate data collection, compare it to operating targets, and data in a way that makes sense to you in order to organize and allow users to share and collaborate on the results instantly withanalyze. It has to host your familiar productivity applications, out picking up the phone, sending an e-mail or walking around. and your custom programs in computer languages that you know.

Taking advantage of this, the refiner created a manufacturing Its inputs and outputs have to be collected and delivered where information system that integrated previously unlinked planning, and when and how you want them. It has to provide access and operations and maintenance information within its plants. The information only to people that you know and trust. system used role-based screens from its collaboration software to drive alignment with operating targets across the plant, from Why is cloud computing important to the HPI— and why now? Because new capabilities that you can use to managers and planners to supervisors and plant operators.

A simple application of business insight: the visibility to transform your business today are out there, waiting for you to the information in a context that made sense to each worker, turn them on. The productivity-improving, collaboration and along with the ability to find and collaborate quickly and easily communication examples that are cited in this article can run with other workers on solutions when targets might be missed, there today. Why wait? helped the refiner improve profit margins while paying out the In the near future, when the CFO enters your office to find IT investment in under one year—all by making better use of out what you’ve done for the business lately, walk him or her the technology on hand. outside and point up. Your answer is in the cloud. HP HYDROCARBON PROCESSING JUNE 2010

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PROCESS INSIGHT Optimizing CO2 Capture, Dehydration and Compression Facilities The removal of CO2 by liquid absorbents is widely implemented in the field of gas processing, chemical production, and coal gasification.

Many power plants are looking at post-combustion CO2 recovery to meet environmental regulations and to produce CO2 for enhanced oil recovery applications. The figure below illustrates actual data of fuel consumption in 2005 and an estimate of energy demand for various fuels from 2010 to 2030.

The world energy demand will likely increase at rates of 10–15% every 10 years. This increase could raise the CO2 emissions by about 50% by 2030 as compared with the current level of CO2 emissions. The industrial countries (North America, Western Europe and OECD Pacific) contribute to this jump in emissions by 70% compared to the rest of the world, and more than 60% of these emissions will come from power generation and industrial sectors.

formulated solvent without implementing any split flow configurations. This is much less than the reported steam usage for the MEA solvent.

The design of a facility to capture 90% of the CO2 from the flue gas of a coal fired power plant is based on the specified flue gas conditions, CO2 product specifications, and constraints. Using the ProMax® process simulation software from Bryan Research & Engineering, CO2 capture units can be designed and optimized for the required CO2 recovery using a variety of amine solvents. The following figure represents a simplified process flow diagram for the proposed CO2 Capture Plant.

Despite the strong recommendations from certain governments, there are very few actual investments in CO2 capture facilities geared toward reducing greenhouse gas emissions mainly because of the high cost of CO2 recovery from flue gas. CO2 capture costs can be minimized, however, by designing an energy efficient gas absorption process. Based on the findings of recent conceptual engineering studies, HTC Purenergy estimated the production cost to be US$ 49/ton CO2 (US$ 54/ tonne CO2) for 90% CO2 recovery of 4 mole% CO2 content in the flue gas of NGCC power plants. A separate study showed the cost for 90% CO2 recovery of 12 mole% CO2 from a coal fired power plant to be US$ 30/ton CO2 (US$ 33/tonne CO2). The cost of CO2 recovery from coal power plant flue gas is substantially less than that of NGCC power plant flue gas due to the higher CO2 content in the feed.

The energy efficiency of a CO2 capture plant depends primarily on the performance of the solvent and optimization of the plant. In traditional flue gas plant designs, MEA was the primary solvent and was limited to 20 wt% to minimize equipment corrosion. Recent developments in controlling corrosion and degradation has allowed an increase in the solvent concentration to about 30 wt% thus decreasing the required circulation and subsequent steam demand. A recent DOE study shows the steam consumption for an existing CO2 plant using 18 wt% MEA (Kerr McGee Process) is 3.45 lb of steam per lb of CO2 for amine regeneration. A modern process that uses 30 wt% MEA is expected to use 1.67 lb of steam per lb of CO2 for amine regeneration. The HTC formulated solvent is a proprietary blend of amines and has a lower steam usage than the conventional MEA solvent. Based on the material and energy balances for the plant designed in the recent study, the reboiler steam consumption is estimated at about 1.47 lb steam/lb CO2 using the proposed

The table below presents the main findings for CO2 capture from the coal fired power plant and the NGCC power plant, each designed to produce about 3307 ton per day (3,000 TPD metric). To produce the same capacity of CO2, only one train with smaller column diameters is required in the case of the coal power plant and two trains with larger column diameters are required in the NGCC Power Plant case. This is mainly due to processing a larger flue gas with lower CO2 content in the NGCC power plant. Consequently, a substantial reduction in the capital and production cost was reported for the coal fired power plant CO2 recovery facility.

For more information about this study, see the full article at www.bre.com/support/technical-articles/gas-treating.aspx.

Bryan Research & Engineering, Inc. P.O. Box 4747 • Bryan, Texas USA • 77805 979-776-5220 • www.bre.com • sales@bre.com Select 113 at www.HydrocarbonProcessing.com/RS

HPIMPACT BILLY THINNES, NEWS EDITOR

BT@HydrocarbonProcessing.com

Large steam-cracking furnace modules arrive in Singapore A significant milestone in the development of ExxonMobil’s largest integrated chemical and refining complex was reached with the arrival of seven world-scale furnace modules at the Singapore facility.

Each furnace module is about 15 stories (50 meters) tall and weighs over 2,000 tons, the equivalent of the combined takeoff weight of five Boeing 747 airplanes (Fig. 1). Applying ExxonMobil’s proprietary furnace technology, the furnaces are part of a feed-flexible steam cracker that will have an ethylene production capacity of one million tons per year.

Engineering work for the furnace modules took 19 months and was carried out in France and Japan and completed in late 2008. A special, multi-functional heavy transport vessel, the RollDock Sun, was selected to ship the furnace modules from Thailand to Singapore (Fig. 2).

“The arrival of these furnaces is a significant milestone for ExxonMobil’s second petrochemical project in Singapore,” said Kwa Chong Seng, chairman and man-

aging director of ExxonMobil Asia Pacific Pte Ltd. “When completed, the integrated manufacturing site here will be well placed to serve the growing markets in the Asia Pacific region.”

ExxonMobil’s second petrochemical project in Singapore includes a world-scale steam cracker and associated derivative units—including new polyethylene, polypropylene and specialty elastomer plants, an aromatics extraction unit and oxo alcohol expansion. This complex will be ExxonMobil’s largest integrated chemical and refining site. The anticipated mechanical completion and startup activities for the new facilities will be phased in beginning in late 2010 through 2011.

—Stephany Romanow

EPA urged to delay judgment on E15 The Auto Alliance, the American Petroleum Institute and the Outdoor Power Equipment Institute urged the US Environmental Protection Agency (EPA) to delay action on the agency’s proposal to allow higher levels of ethanol in gasoline. Higher levels of ethanol have not

been proven safe or effective according to industry projections based on preliminary results of testing introduced at a meeting of the Mid-Level Ethanol Blends Research Coordination Group.

Dave McCurdy, president and CEO of the Alliance of Automobile Manufacturers; Jack Gerard, president and CEO of the American Petroleum Institute; and Kris Kiser, executive vice president of the Outdoor Power Equipment Institute issued this statement: “As the EPA proceeds with important decisions about ethanol and biofuel blend rates, it is imperative that those decisions be made with the end-user market in mind. These decisions will have real world impacts and we urge the EPA to refrain from setting a premature deadline that ignores reliable, scientific data about the effects of higher ethanol blends on emissions, durability and consumer safety.

We remain committed to finding the right market solutions for renewable fuels and look forward to continuing our work with the EPA on this matter.”

In addition to government funds, the auto and oil industries have spent more than $6 million over the last two years testing engine performance and durabil-

FIG. 2 FIG. 1

Completed furnace modules were shipped to ExxonMobil’s Singapore petrochemical complex on a specially designed heavy transport vessel.

Placement of a furnace module, which is 15 stories tall and weighs more than 2,000 tons.

HYDROCARBON PROCESSING JUNE 2010

I 19

HPIMPACT ity of higher ethanol fuels, as well as testing storage and dispensing of fuels with 15% ethanol (E15). At present, fuels are allowed by the EPA to contain only up to 10% ethanol (E10).

Approval of E15 should wait until testing is complete. This testing looks at the potential for vehicle engine and fuel system component damage when operating on E15. “The impacts of higher ethanol blends will fall on consumers, who will be

ill-prepared to determine the right fuel for their car, lawn equipment, boat or motorcycle,” said Al Jessel, senior fuels policy advisor for Chevron. “The EPA should delay changing the gasoline mix in this country until research into all aspects of vehicle and engine performance is complete.”

“Testing for engine and vehicle compatibility and environmental issues is scheduled for completion in 2011,” said Coleman Jones, biofuel implementation

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manager for General Motors. “There’s no need for precipitous action when the scientific results are so close at hand.”

—Stephany Romanow

Energy future must include mix of renewables Meeting ever-rising demand with hydrocarbon reserves that are becoming more difficult and expensive to exploit will require a makeover of the oil and gas industry over the next 40 years, a panel of operators and government officials told a general session at the Offshore Technology Conference (OTC), held in Houston the first week of May. With even the most conservative mainstream forecasts showing energy consumption rising by more than 40% over the next 20 years, the pressure on identifying and developing new conventional and unconventional fossil fuel reserves will be intense, the panelists said. At the same time, it also is incumbent on traditional oil and gas companies to help further the economical development of renewable resources.

Like many of this year’s OTC technical and general sessions, the discussion quickly turned from the advertised topic of what is the right global energy mix to the ongoing Deepwater Horizon tragedy.

“I was going to look at the future and discuss the priority for technology and science, but as I stand before you, my thoughts are on the Gulf of Mexico and the loss of our colleagues and the environmental challenge we are confronting,” said panel moderator Ahmed Hashami, BP vice president of technology management and finance. “It’s a challenging time for us and the industry. We are using the full resources of BP and the industry to deal with this challenge and to take every measure to ensure it never happens again.”

Without mentioning the Gulf of Mexico catastrophe, a US Department of Energy (DOE) official followed, saying the federal government will continue its research and development efforts for all energy sources, including oil and gas. She echoed the current administration’s battle cry that any energy mix must include alternative and clean technologies.

“Environmental stewardship and increased use of fossil fuels need not be mutually exclusive,” said Deputy Assistant DOE Secretary Phyllis Yoshida. “We in the Department of Energy believe firmly

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We put the best hands in the business to work on your project.

When it comes to turnarounds, no one can beat the loyalty, dedication to quality and pure craftsmanship of our “hands” in the field. Through the years AltairStrickland has developed a following of skilled workers. Many have worked on the same projects together for a decade or more. Their familiarity with one another’s abilities, talents and experience means a more cohesive effort and efficient project execution for you.

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Call us for your next turnaround, revamp, upgrade, repair or emergency; we’ll put the best hands (and minds) in the business to work for you.

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HPIMPACT that technology and science are critical to maintain and enhance our standard of living. The decisions we make today can impact our world for a generation. Right now, the US is importing 60% of its oil, so, to reduce uncertainties, we must make investments for new energy sources here, including fossil fuels.”

Even though Saudi Aramco remains a powerhouse oil and gas producer, its executive director said real-world realities demand that it, too, look at renewable energy sources. “Energy is too important to have a future clouded by anxiety. Decisions cannot be made today based on over-heated emotions,” said Mohammed Al-Qahtani, adding that the Kingdom alone still has more than 4.7 trillion bbl of proven oil reserves and well over 275 tcf of known gas.

“By all reliable standards of measurement, the world has enough reliable reserves for several more decades,” Mr. AlQahtani said. “However, we as an industry will be more diverse than we’ve been over the last 100 years. We must constantly assess how we do our business. We need to increase the momentum for gas and we must maintain and increase our investments in renewable resources and environmental efficiency.”

Chevron President Ali Moshiri agreed, saying that all avenues must be on the table and this includes continued research and development of new oil and gas reserves.

“The proper energy mix between fossil fuels and alternative sources has no right or wrong answer. No one size fits all, but it has to be decided by the marketplace and not

government. Our industry has delivered reliable supplies for 100 years, often during some very turbulent times,” he said.

Acknowledging that energy use is expected to rise by as much as 40% over the next 20 years, Matthias Bichsel, director of projects and technology for Royal Dutch Shell, said that oil and gas must play the predominate role. “There is going to be even more pressure on oil and gas reserves, but, at the same time, we have

For additional coverage of OTC, visit the HPInformer blog (www.hydrocarbonprocessing.com/hpinformer). Hydrocarbon Processing’s sister publication World Oil published the official show daily newspaper for the event (a total of four editions).

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Ali Moshiri, president of Chevron, said that there is no right or wrong answer to the question of a proper energy mix.

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© 2010 Thermo Fisher Scientific Inc. All rights reserved. Copyrights in and to the Matches and Blowtorch photographs are owned by a third party and licensed for limited use only to Thermo Fisher Scientific by maxx images and SuperStock.

Imagine one gas analyzer that does the work of ten.

Introducing Thermo Scientific Prima PRO, the online gas analyzer that uses high-speed mass spectrometry to increase yield, conserve energy, and enhance process safety.

In fact, a single Prima PRO can do the work of ten gas chromatographs to dramatically lower your cost of ownership and raise the bar on product quality.

This powerful yet easy-to-use, simple-to-maintain system offers automatic tuning and calibration and produces higher resolution for a number of process variables. So you have a better understanding of process dynamics and can implement more effective operating procedures. Plus, Prima PRO is built to deliver 99.9% uptime—making it a safe, reliable choice.

Thermo Scientific Prima PRO— tap the power of online process mass spectrometry to optimize plant productivity and accelerate profits.

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Moving science forward Select 115 at www.HydrocarbonProcessing.com/RS

HPINNOVATIONS SELECTED BY HYDROCARBON PROCESSING EDITORS editorial@gulfpub.com

Agreement made for MSAR Quadrise Fuels International plc (QFI), the producer of oil-in-water emulsion fuel as a low-cost substitute for heavy fuel oil for use in power plants and industrial diesel engines, just signed a joint development agreement with A.P. Moller–Maersk. Over the next 12 months, the agreement jointly investigates and develops a marine version of its multiphase superfine atomized residue (MSAR) emulsion fuel. Maersk is the world’s largest container shipping organization and the largest single purchaser of bunker fuel.

The development program will capitalize on QFI’s skills in emulsion fuel application and use in large diesel engines. Maersk will contribute their vast shipping industry experience and provide unsurpassed technical expertise relating to naval architecture, machinery systems, fuels, exhaust gas emissions and technical ship management.

Maersk vessels provide excellent testing platforms for new technologies.

In the event the development program is successful, Maersk and QFI will seek to expediently implement the technology. Bill Howe, chief executive officer of QFI stated, “This cooperative development is a very exciting opportunity for QFI. The diesel engine application of MSAR represents a market sector with limited inter-fuels competition. Bunker fuel, driven by expansion in international trade, is a long-term growth market and comprises 150 million tons annually, approximately 30% of the global heavy fuel oil market. Bunker fuel represents a high proportion of shipping industry operating costs and there is great potential for MSAR if the technology can be successfully adapted to the marine application.”

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Evolutionary facepiece seal technology Scott Health & Safety announced the launch of its SureSeal system, the most dynamic facepiece seal available today for users of respiratory protection products.

Developed with numerous innovations such as an enhanced sealing surface and practical personal protective equipment (PPE) interfaces, the Scott SureSeal system combines the preference for greater user comfort with

the need for superior fit. Since it is integrated with the Scott proven AV-3000 facepiece, users experience an expansive field of vision and exceptional voice intelligibility.

“With this new design, we have taken our industry-leading Scott AV-3000 facepiece to the next level in comfort and fit,” said Marty Lorkowski, Scott’s global industrial marketing manager. “We did it by taking the best from our current facepiece and combining it with a completely redesigned seal technology to develop our most innovative facepiece yet.”

The SureSeal system utilizes a U-shaped seal that maintains a continuous circumferential seal around the user’s face (Fig. 1). This patent-pending Scott-developed design, called reverse reflex, allows the SureSeal system to flex with every movement of the facepiece, resulting in a higher fit factor and a more comfortable seal for the user.

Additionally, a fifth strap improves sealing capabilities by adding another point of connection to the smaller Kevlar head harness.

The tightening buckles are thoughtfully placed so they won’t torque the seal and cause leaks if over-tightened.

A primary feature of the SureSeal system includes a deeper, lower-profile seal that offers a more robust interface with protective suits, hoods, shrouds and helmets. Additionally, a smaller Kevlar head harness eliminates bunching and discomfort under helmets and hard hats. The fifth strap is strategically positioned so that it won’t interfere with the head suspension of protective helmets. For proper and quick donning, a reflective positioning tab is located on the back of the head harness.

As with all of Scott’s AV-series facepieces, the AV-3000 facepiece with the SureSeal system continues Scott’s top-down convertibility design concept, which allows for a As HP editors, we hear about new products, patents, software, processes, services, etc., that are true industry innovations—a cut above the typical product offerings. This section enables us to highlight these significant developments.

For more information from these companies, please go to our Website at www.

HydrocarbonProcessing.com/rs and select the reader service number.

single facepiece to be used for all NFPA/ CBRN/NIOSH respiratory applications.

According to Mr. Lorkowski, “It was vitally important to maintain top-down convertibility when designing the SureSeal system.

We know that our current customer relies on this design concept to reduce costs associated with larger equipment inventories and fit testing for multiple facepieces.”

The SureSeal system is available on all new AV-3000 facepieces and can be retrofitted to AV-3000 facepieces currently in use.

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Software that wirelessly tracks supplies and improves safety Savi Technology, a Lockheed Martin company announced that its SmartChain solution can now be leveraged by the oil and gas industry to improve supply chain and work-related activities, reduce downtime and provide a consolidated view of real-time data to make better decisions.

The SmartChain release automates information on the status and integrity of assets and products, as well as the safety of personnel throughout the oil and gas value chain, including in exploration and production, refining and distribution operations.

SmartChain is enterprise-level software that optimizes operational readiness, improves asset utilization, and minimizes unplanned events by providing actionable

FIG. 1

SureSeal facepiece.

HYDROCARBON PROCESSING JUNE 2010

I 25

HPINNOVATIONS information on the location and status of assets such as cargo carrying-units, rail cars, tools and drill pipe. SmartChain also monitors the integrity of mission-critical equipment, such as pipeline, and the status of products and personnel in refineries.

“Not having the right tool at an offshore platform or oil rig can actually shut down the whole process for a day or two, costing tens of millions of dollars,” said Dr. Ben Zhogi, a professor at Texas A&M University

and director of the radio-frequency identification (RFID) technologies oil and gas consortium. “RFID technology solutions can relieve new challenges faced in the petroleum industry by helping companies more effectively utilize assets at a time of high oil demand worldwide and constrained capacities, and to detect potential system failures and locate personnel in times of crisis.”

“Already proven in the world’s largest and most complex supply chains, our

SmartChain solution can revolutionize the efficiency and effectiveness of oil and gas operations by providing true end-to-end visibility,” said David Stephens, chief executive officer of Savi Technology. “SmartChain for the petroleum industry draws on two decades of experience providing complete wireless supply chain solutions, coupled with domain and technology expertise from our partners.”

The highly adaptable SmartChain enterprise platform and asset management suite, already proven in the chemical, aerospace and defense sectors, leverages wireless tracking technologies to automatically locate and manage equipment and supplies required in petroleum exploration and production.

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Heat and power solution incorporating VOC destruction

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Do you have flows up to 1,400 US GPM (320 m3/hr), RS Series heads up to 3,400 feet (1,000 m), pressures up to 1,500 psig (100 bar), temperatures from 20˚F to 300˚F (-30˚C to 149˚C), and speeds up to 3,500 RPM? Then you need Carver Pump RS Series muscle!

Designed for moderate to high pressure pumping applications, the RS is available in five basic sizes with overall performance to 1,000HP. As a standard, with a product lubricated radial sleeve bearing and two matched angular contact ball bearings for thrust, it only takes a mechanical seal on the low pressure, suction side to seal the pump. Optional features include ball bearings on both ends with an outboard mechanical seal, various seal flushing arrangements and bearing frame cooling. These features make the RS ideally suited for Industrial and Process applications including Pressure Boost Systems, Boiler Feed, Reverse Osmosis, Desalination and Mine Dewatering. Whatever your application, let us build the muscle you need!

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Siemens Energy recently entered into an agreement with Environment and Power Systems International to work with US industrial customers on combined heat and power project solutions incorporating the destruction of unwanted volatile organic compounds (VOC). The sensible use of VOCs in the SGT-300 cogeneration system designed for industry is a pivotal wasteto-energy technology for major sources of environmentally regulated VOCs in the manufacturing, petrochemical, synthetic and organic manufacturing industries.

The single-shaft SGT-300 gas turbine combines advanced technology with a rugged industrial design. It has a power output of 7.9 MW and is available as a factory-assembled package. The SGT-300 is designed to operate on a wide range of gaseous and liquid fuels and is equipped with a dual-fuel dry low emissions combustion system, meeting the most stringent legislation for NOx. This state-of-the-art combustor design is ideal for ingesting vaporized and gaseous VOCs that are thermally oxidized into the end-products CO2 and water.

The waste VOC emissions can be captured, conveyed and ingested into the turbine’s air intake. Not only are the harmful hydrocarbons destroyed but the VOC is utilized in the combustion chamber as a supplemental fuel in addition to the natural gas fuel that is directly injected into the combustor to fuel engine operation. This technology has been successfully demonstrated in an industrial application in the Netherlands.

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HPIN CONSTRUCTION BILLY THINNES, NEWS EDITOR BT@HydrocarbonProcessing.com

North America Dominion and Exterran Holdings, Inc., will team up to provide increased natural gas processing in the Appalachian Basin of the US. Under the agreement, two natural gas processing plants will be designed, built and installed by Exterran. One will be an 8 million-cf/d-capacity plant in Carlisle, Ohio, and the other a 10 million-cf/dcapacity plant in Schultz, West Virginia.

The two plants will be owned and operated

by Exterran. Construction is expected to begin immediately. Commercial operation of the Carlisle plant is anticipated in January 2011 and commercial operation of the Schultz plant is anticipated in May 2011.

ONEOK Partners, LP, will invest approximately $405 million to $470 million between now and the end of 2011 for projects in the Bakken shale in the Williston Basin in North Dakota and the Woodford shale in Oklahoma. These investments include construction of a new 100 millioncf/d natural gas processing facility in eastern McKenzie County, North Dakota, and related expansions that are estimated to cost between $150 million and $210 million and will double the partnership’s natural gas processing capacity in the Williston Basin. Completion is expected in the fourth quarter of 2011.

Samsung Engineering has a $220 million ultra-low-sulfur diesel unit plant project order with the Petroleum Co. of Trinidad & Tobago (PETROTRIN). The plant will be constructed in the Pointea-Pierre refinery complex and will be producing 40,000 bpd of ultra-low-sulfur diesel, in compliance with EURO-V standards. This project will be executed on a lump-sum turnkey basis that includes all engineering, procurement, construction and commissioning and its expected completion is in March 2012.

Enterprise Products Operating LLC has a six-year agreement to provide Anadarko Energy Services Co. with natural gas liquid (NGL) fractionation services.

Under the terms of the contract, Enterprise will make up to 62,000 bpd of firm NGL fractionation capacity available at

the partnership’s fractionation complex in Mont Belvieu, Texas, beginning September 1, 2010.

Europe Foster Wheeler AG’s Global Engineering and Construction Group has been awarded contracts by PETROM S.A. for a major refinery modernization project being implemented at the Petrobrazi refinery in Ploesti, Romania. Foster Wheeler will provide front-end engineering design (FEED) modification services and engineering, procurement and construction management (EPCm) for the revamp of an atmospheric/vacuum distillation unit at the refinery. In addition, the company has been awarded an EPCm contract for a new amine unit and the FEED for the revamp of a delayed coking unit.

Jacobs Engineering Group Inc. has a three-year extension to its existing €200 million per year integrated services contract (ISC) for Shell Onegas. The work is being executed onshore at gas plants in Bacton, UK, and Den Helder, The Netherlands, and offshore on 54 assets in the TREND ANALYSIS FORECASTING Southern North Sea. Hydrocarbon Processing maintains an extensive database of historical HPI projGAZPROM LLC ect information. Komplektatsiya Current project activity is published times a yearprogram in the HPIto has launchedthree an investment Construction Boxscore.

When a project increase the octane number of the existing is completed, it is removed from current gasoline production and also the profitlistings and retained in a database. The ability of its refinery located in Astrakhan, database is a 35-year compilation of projects by type, operating company, licenRussia.

The refinery was first put in service engineering/constructor, location, insor, 1985.

The contract for this projectetc. was Many companies use the historical data for signed with MAVEG Industrieausrüstuntrending or sales forecasting.

gen GmbH. Lurgi GmbH and MAVEG The historical information is available in were selected to supply and comma-delimited or Excel®engineering and can be cusprocurement this The project tom sorted toservices suit yourfor needs. cost and of thehandle sort depends on thehydrotreater size and complexwill a naphtha and a of the sort you request and whether a City unit to be integrated 5/C6 isomerization customized program must be written. You into refinery.

Startuprequest of the overall canthe focus on a narrow such as plant the is history scheduled 2012. type of project or of afor particular you can obtain the entire 35-year Boxscore database, or portions thereof.

The partners in the Gassled joint venSimply send aGassco’s clear description of the datato ture approved recommendation you need and you will receive a prompt build new terminal rather than upgrading cost aquotation.

Contact: the existing facility to extend the lifetime Lee Nichols for the NorseaP. Gas Terminal (NGT) in O. Box 2608 Texas, 77252-2608 Emden, Houston, Germany.

C&BI Lummus B.V. Fax: 713-525-4626 has been awarded the pre-engineering cone-mail: Lee.Nichols@gulfpub.com. tract for the project.

Chicago Bridge & Iron Co. N.V. has an agreement with CB&I Lummus for concept development services for the Yamal LNG integrated project. This contract is scheduled for completion in the first half of 2011. The project consists of the production, treatment, transportation, liquefaction and shipping of natural gas and natural gas liquids from the South Tambey field on the Yamal Peninsula in Northwestern Siberia, Russia. CB&I’s project scope includes concept development of the 15–16 million-tpy LNG liquefaction plant, including LNG storage and loading facilities.

Middle East Qatargas has awarded Fluor the engineering, procurement and construction management contract for the Jetty boil-off gas recovery project. This project at Ras Laffan Industrial City, Qatar, will minimize LNG boil-off gas flaring at LNG berths by making productive use of the gas that boils off during loading of LNG carriers at the Ras Laffan Port. The planned completion date for this project is at the end of 2013 or early 2014.

Celanese Corp. and Saudi Basic Industries Corp. (SABIC) announced that their National Methanol Co. (Ibn Sina) joint venture will construct a 50,000 ton polyacetal TREND ANALYSIS FORECASTING Hydrocarbon Processing maintains an extensive database of historical HPI project information. The Boxscore Database is a 35-year compilation of projects by type, operating company, licensor, engineering/constructor, location, etc. Many companies use the historical data for trending or sales forecasting.

The historical information is available in comma-delimited or Excel® and can be custom sorted to suit your needs. The cost of the sort depends on the size and complexity of the sort you request and whether a customized program must be written. You can focus on a narrow request such as the history of a particular type of project or you can obtain the entire 35-year Boxscore database, or portions thereof.

Simply send a clear description of the data you need and you will receive a prompt cost quotation. Contact: Lee Nichols P. O. Box 2608, Houston, Texas, 77252-2608 Fax: 713-525-4626 e-mail: Lee.Nichols@gulfpub.com HYDROCARBON PROCESSING JUNE 2010

I 29

HPIN CONSTRUCTION (POM) production facility in Saudi Arabia.

Engineering and construction of the facility is expected to begin later this year. Construction of the facility is part of an extension of the Ibn Sina joint venture, which will now run through 2032. Ibn Sina produces methanol, a key feedstock for POM production, as well as methyl tertiary-butyl ether.

Total recently started up an olefin cracker based on ethane in Ras Laffan,

Qatar. With a production capacity of 1.3 million tpy of ethylene, the Ras Laffan olefin cracker will feed the new Qatofin polyethylene plant. The ethane feedstock used in the Ras Laffan cracker comes from the North Field. The natural gas (methane) is treated for export in the liquefaction plants also based in Ras Laffan and the associated ethane produced will be valorized by the Ras Laffan cracker as raw material for the petrochemical industry.

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I JUNE 2010 HYDROCARBON PROCESSING

Asia-Pacific Royal Vopak recently opened an import and distribution terminal for oil products in Jakarta, Indonesia. The newly built terminal in Tanjung Priok will support the storage and distribution needs of the Indonesian and international oil industry. The terminal has a storage capacity of 250,800 cubic meters and enables the distribution of oil products in the greater Jakarta area.

The terminal’s storage capacity can be expanded to 450,000 cubic meters in the future depending on market demand.

CB&I’s ethylene cracker project in Singapore has started up, producing onspecification ethylene and propylene. The cracker complex is owned and operated by Shell Eastern Petroleum Ltd. and is a part of the Shell Eastern petrochemicals complex project in Singapore, which comprises modifications to the existing Bukom refinery as well as the building of a new mono-ethylene glycol plant on Jurong Island. The 800,000tpy cracker increases Singapore’s ethylene capacity by 40% while also producing 450,000 tpy of propylene, 230,000 tpy of benzene and 155,000 tpy of butadiene.

ConocoPhillips has an agreement with POSCO to use ConocoPhillips’ technology in POSCO’s Gwangyang, Korea, coalto-substitute natural gas (SNG) project.

The project will allow POSCO to produce 500,000 metric tpy of SNG from the gasification of approximately 1.8 million tons of coal. Preliminary design work began in 2008 and site preparation is now underway. Under the agreement, ConocoPhillips will provide process engineering design and technical support relating to the gasification technology process block of the SNG facility.

Alfa Laval has an order from an Indian refinery for a heat exchanger. The order value is about 95 MSEK and delivery is scheduled for 2011. The plate heat exchanger, with a height of 25 meters, will be used in a catalytic reforming unit for the production of gasoline in an Indian refinery.

Larsen & Toubro has an order valued at Rs. 1400 crore from Indian Oil Corp.

Ltd. for a 4.17 million-metric tpy FCC reactor regenerator project of a grassroots fuel refinery at Paradip, Orissa, India. The scope of work includes residual and detailed engineering, procurement, supply, manufacture, fabrication, construction, installation and performance test run. HP

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HPI CONSTRUCTION BOXSCORE UPDATE Company

City

Plant Site

Project

Capacity Unit Cost Status Yr Cmpl Licensor

Engineering

Constructor

UNITED STATES Illinois Indiana Louisiana North Dakota Ohio Texas West Virginia

Abengoa Bioenergy Abengoa Bioenergy Shintech Inc. Oneok Corp.

Dominion/Exterran JV Anadarko Dominion/Exterran JV

Madison Mt Vernon Plaquemine McKenzie Co Carlisle Mont Belvieu Schultz

Madison Mt Vernon Plaquemine Garden Creek Carlisle Mont Belvieu Schultz

Ethanol Ethanol VCM (2) Natural Gas Plant Natural Gas Plant NGL Fractionation Natural Gas Plant

BY

88 88 800 100 8 102 10

MMgpy 200 MMgpy 200 Mtpy 1064 MMcfd 200 MMcfd Mbpd MMcfd

C C U F U P U

2010 2010 2011 2011 2011 2011 2011

P C E

2013 2010 2012

P F S F F F F P P P E

2011 2011 2015 2011 2011 2011 2011 2011 2011 2011 2013

U S S S S S S U U U C P E F F

2012 2012 2012 2012 2012 2012 2012 2011 2013 2011 2010 2012 2012 2011 2012

Vogelbusch Vogelbusch

Enterprise Products

LATIN AMERICA Brazil Brazil Trinidad

Petrobras Hexion Specialty Chemicals Petrotrin

Para Para Biodiesel Rio Grande do Sul Rio Grande do Sul Urea-Formaldehyde Pointe-a-Pierre Pointe-a-Pierre Diesel, ULSD (1)

120 MMl/y 450 Mm-tpy 40 Mbpd

Dioki d.d.

Gassco Petrobras/Galp Energia JV Petrom Petrom Petrom Petrom Yamal LNG LLC Yamal LNG LLC Yamal LNG LLC Turkmenhimiya

Krk Emden Porto Ploesti Ploesti Ploesti Ploesti Yamal Yamal Yamal Mary City

Krk Emden Porto Ploesti Ploesti Ploesti Ploesti Yamal Yamal Yamal Mary City

Polyvinyl Chloride (PVC) Terminal, Gas Biodiesel Amine Unit Distillation, ADU/VDU Hydrocracker Hydrogen LNG Liquefaction Plant LNG Storage Utilities Urea

120 Mtpy None 250 Mtpy None 1.5 MMtpy None None 15 MMtpy None None 636 Mtpy

Jiujiang Shaanxi Shaanxi Shaanxi Shaanxi Shaanxi Shaanxi Tianjin Wuhan Wuxi Zhenhai Panipat Paradip Paradip Gwangyang

Jiujiang Shaanxi Shaanxi Shaanxi Shaanxi Shaanxi Shaanxi Tianjin Wuhan Wuxi Zhenhai Panipat Paradip Paradip Gwangyang

Petrochemical Complex BY 800 Mtpy Alcohol Iso Nonyl 200 Mtpy Ethylene Propylene Rubber 200 Mtpy Methanol 1.8 MMtpy Methanol-to-Olefins (MTO) 600 Mtpy Polyethylene 450 Mtpy Polypropylene 250 Mtpy Terminal 950 Mm3 Ethylene 800 Mtpy Plastics (2) TO 60 Mm-tpy Ethylene 1 MMtpy SBR Elastomers 120 Mtpy Cracker, FCC Reactor/Regenerator 4.17 MMtpy Storage, Tank TO 30 MMtpy Coal to SNG Plant 500 Mtpy

50 60 220

Lummus Technology

Samsung Eng

EUROPE Croatia Germany Portugal Romania Romania Romania Romania Russian Federation Russian Federation Russian Federation Turkmenistan

EX

10 240 1332 1332 1332 1332

662

CB&I Lummus FW FW CB&I CB&I CB&I Sojitz Corp|Kawasaki Plant Systems

MWKL

ASIA/PACIFIC China China China China China China China China China China China India India India South Korea

Sinopec Jiujiang Petrochemical Yanchang Petroleum Group Yanchang Petroleum Group Shaanxi Yanchang Petroleum Yanchang Petroleum Group Yanchang Petroleum Group Yanchang Petroleum Group Sinochem Sinopec\SK Energy JV Lanxess Sinopec Intl Co IOCL/TSRC/Marubeni JV IOCL Paradip Port Trust POSCO STEEL

3200 3200 3200 3200 3200 3200 51

203 317 112

Lanxess Indian Oil Essar ConocoPhillips

L&T

See http://www.HydrocarbonProcessing.com/bxsymbols for licensor, engineering and construction companies’ abbreviations, along with the complete update of the HPI Construction Boxscore.

BOXSCORE DATABASE

ONLINE

THE GLOBAL SOURCE FOR TRACKING HPI CONSTRUCTION ACTIVITY For more than 50 years, Hydrocarbon Processing magazine remains the only source that collects and maintains data specifically for the HPI community, publishing up-to-the-minute construction projects from around the globe with our online product, Boxscore Database. Updated weekly, our database helps engineers, contractors and marketing personnel identify active HPI construction projects around the world to: • Generate leads • Market research • Track trend analysis • And, decide future budget planning.

Now, we’ve made our best product even better!

Enhancements include: • Exporting your search results to Excel so you can compile your research • Delivering the latest updated projects directly to your inbox each week • Designing customized construction reports for your company using our 50 years of archived projects.

For a Free 2 -Week Trial, contact Lee Nichols at +1 (713) 525-4626, Lee.Nichols@GulfPub.com, or visit www.ConstructionBoxscore.com

32

I JUNE 2010 HYDROCARBON PROCESSING

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So you can install more devices on a process plant bus network. Or use them in remote locations with solar/battery sources. And unlike integrated valves, you can choose from a wide range of easily available models, with larger orifices to handle higher flows without clogging. Many ASCO low-power solenoid valves come with the ASCO Today same-day shipping program for the fastest delivery on the planet.

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PROCESS/PLANT OPTIMIZATION

SPECIALREPORT

How will carbon emissions regulations revise energy conservation economics?

Including the cost of carbon in refinery project economics has the potential to convert previously marginal energy projects into more attractive options I. M. GLASGOW, S. POLCAR, E. DAVIS, T. NGUYEN, J. PRICE, C. STUECHELI and R. E. PALMER, Mustang Engineering, L.P., Houston, Texas

A

program directed at reducing greenhouse gas (GHG) emissions is gaining higher interest by US public policy makers. This is evident by the House of Representatives passing The American Energy and Security Act (H.R. 2454) in June 2009. The US EPA is actively evaluating avenues in which the Clean Air Act can be used to reduce GHG emissions based on its determination that these substances (GHGs) are an endangerment to human health and welfare.

In H.R. 2454, individual refiners would be responsible for GHG emissions from their manufacturing operations plus the emissions from the combustion of the fuels sold by the refinery.

These total emissions represent about 35% of the total US GHG inventory. However, in H.R. 2454, the refining sector is given 2% of the available emission allowances per year until 2025. Result: Refiners will be required to purchase or to find offsets for over 90% of their regulated GHG emissions.

In this potential carbon-constrained economy, significant incentives to implement energy conservation projects that are marginal or uneconomic based only on the value of fuel savings will be investigated. In this article, several case studies will be presented in which refinery energy conservation options will be considered with the economics for GHG emissions reductions. A sensitivity analysis on the value of the GHG allowances will be included in the economic evaluation.

GHGs that are actually emitted during the refinery manufacturing process and subsequent combustion of fuels by the transportation sector. As a result, refiners will incur significant expenses to reduce manufacturing GHG emissions and to purchase additional allowances required to comply with the cap-and-trade rules.

Other regulatory approaches of stationary sources are being evaluated by the EPA. This consists of mandatory reporting of GHG emissions, and a proposed rulemaking that will require best available control technology (BACT) of GHG emissions for Title V and PSD Permits. In the absence of federal legislation, regional cap-and-trade markets have been established, as shown in Fig. 1.

Case studies. A 150,000-bpsd (150-Mbpsd) notional refinery will be used to demonstrate the magnitude of GHG emissions and potential energy efficiency projects. Arab Medium crude was

Background. GHGs are generally

defined as the total of six compounds, the main contribution being from energyrelated carbon dioxide (CO2) produced by combusting fossil fuels.

Of the annual GHG allowances available in the proposed Waxman-Markey cap-andtrade bill, only 2% of the available emissions will be allocated to the refining sector.

This compares to about 35% of the total US

FIG. 1

Regional climate initiatives.

HYDROCARBON PROCESSING JUNE 2010

I 35

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION

selected as the basis. Hydrogen is supplied from an onsite generation facility and a catalytic reformer (CCR). Gasoil upgrading is via a cat-feed hydrotreater and fluid catalytic cracker (FCC). Resid upgrading is achieved via a delayed coker. It is also assumed that flare-gas recovery has already been installed. This notional refinery produces these products: • 8.4 Mbpsd liquefied petroleum gas (LPG) • 78.8 Mbpsd gasoline • 14.8 Mbpsd jet fuel • 43.0 Mbpsd ultra-low-sulfur diesel (ULSD) • 3.4 Mbpsd No. 6 fuel oil • 1,260 tpd coke.

Fig. 2 shows the GHG emissions from the notional refinery.

Emissions from the combustion of transportation fuels sold by

the refinery make up 80% of the potentially regulated emissions. Emissions generated from co-products (coke and LPG) are another 10%. Only 10% of the emissions result from the refinery manufacturing process.

Incremental fuel for the refinery is assumed to be supplied by natural gas. It is assumed that GHG changes in electricity consumption related to offsets are compared to a coal-fired power plant. The refinery is assumed to operate 8,400 hr/yr.

Energy prices are assumed to be $5.50/MMBtu for natural gas on an LHV basis, $7.70/1,000 lb of steam and $0.07/kWh of electricity. The target internal rate of return (IRR) for plant energy conservation projects is assumed to be 20%. The energy conservation projects under review are: • Revamp crude preheat exchange train to increase the crude heater inlet temperature • Add combustion air preheat to fired Distribution of refinery emissions GHG manufacturing emissions heaters Misc other LPG/coke • Replace the vacuum tower steam ejec0.4% 9% tors with a liquid-ring compressor FCC, reforming Plant catalyst • Revamp the diesel hydrodesulfuri10% 35.6% Heaters, boilers zation (HDS) hot feed with additional 50.1% steam generation • Install power recovery from FCC regenerator hot flue gas.

A review of the economics over a range of GHG emission prices is provided for each case.

Gasoline, diesel, jet, marine No. 6 81%

FIG. 2

H2 plant 13.6%

Flares 0.3%

Crude heater inlet temperature.

Notional refinery GHG emissions.

Top PA LVGO

Crude 150,000 bpd 60°F

Top PA

Product 330°F

LVGO

Kero

To cooler 235°F

To cooler 235°F

HVGO

450°F

Diesel Top PA return

To cooler 265°F

Desalted crude

Kero prod Kero PA 250°F

Kero PA

Kero PA return

260°F Desalter AGO prod Diesel prod Diesel PA HVGO

Diesel PA

AGO Product 306°F

AGO PA VTB HVGO AGO PA Steam return 48,000 lb/hr

415°F Diesel PA return

36

Current crude preheat exchange network.

I JUNE 2010 HYDROCARBON PROCESSING

71 MMBtu/hr 510°F

Quench to vac twr

VTB to coker BFW HVGO PA

FIG. 3

AGO PA Vac twr bottoms Crude to heater 475°F

For this case, it is assumed that the crude unit was originally sized for 110 Mbpsd of feed and achieved a heater inlet temperature of 500°F. Over time, the crude rate has increased to 150 Mbpsd by minor revamps with no changes to the preheat exchange network. The existing preheat train uses four pumparounds (PAs) from the crude tower and two PAs from the vacuum tower.

All side-cut products from the crude and vacuum towers are cooled against the crude feed including the vacuum tower bottoms (VTB). Fig. 3 is a summary flow sheet of the current 150 Mbpd notional crude unit.

The project’s goal is to increase this crude heater inlet temperature. The revamp design attempts to target thermal pinch limits to obtain maximum available duty by using the existing available hot streams.

Care is needed for any changes in wasteheat steam production and subsequent changes in fuel consumption/GHG emissions at the boiler house.

Fig. 5 shows a revised heat-exchange configuration that can increase the charge heater temperature to 515°F. New exchangers were sized with appropriate metallurgy, and it was assumed that existing plot space was available for the new equipment. Since a parallel path of exchangers was selected, no changes were required for the pumps.

The capital cost for this revamp was esti-

PROCESS/PLANT OPTIMIZATION

Combustion air preheat. Traditionally, most refinery heaters

$15/metric ton would extend to heaters with an absorbed duty down to 100 MMBtu/hr.

Vacuum tower steam ejectors. The notional refinery uses

a three-stage ejector system (Fig. 8). Each ejector is coupled with an inter/after condenser with all condensed materials sent to a Crude preheat revamp 30 25 20 IRR, %

mated to be $19.5 million. Table 1 summarizes the process data for the case study.

The IRR for GHG emissions prices from $0/metric ton to $100/metric ton is summarized in Fig. 4. A debit in steam generated in the coker due to the decreased feed temperature from the vacuum tower is accounted for in the economics. The result shows that such a revamp is marginal based on fuel savings alone.

However, with the additional cost that will result from GHG emissions in a proposed cap-and-trade system, the project economics are significantly more attractive. A GHG cost of approximately $50/metric ton is required to achieve the 20% IRR target.

Installing welded-plate heat exchangers or other newer exchanger technologies could achieve an even greater increase in the crude heater inlet temperature by achieving tighter thermal pinch.

SPECIALREPORT

15 10

440°F

443°F

445°F

5 have a natural draft design with the process fluid flowing through both the convection and radiant sections. For larger heaters, the 0 20 40 0 60 80 100 additional convection surface may be used for steam generation to GHG cost, $/metric ton obtain efficiencies of 88%–89%. Conversely, for large heaters where higher efficiencies can be justified, or if incremental steam generaFIG. 4 Crude preheat exchange revamp economics with cost of GHG emissions included. tion capability has little or no value, combustion air preheaters have been applied. Several past optimization studies have indicated that when fuel gas value is relatively high, air preheat is economically 27.2 MMBtu/hr 21.6 MMBtu/hr attractive for heaters above 75 MMBtu/hr 5,400 ft2 2 5,000 ft Kero Kero PA absorbed duty.

Top PA Smaller heaters with older designs may LVGO have few convection rows with efficiencies of about 82%. With the additional cost of (Add 1 Kero prod To cooler shell Product GHG emissions, these heaters may become 195°F 2 Kero PA LVGO 4,000 ft ) 330°F good candidates for air preheaters. It is Crude Top Kero assumed that adding air preheat will increase 150,000 bpd HVGO 275°F PA PA 60°F the heater efficiency from 82% to 92%.

245°F Cost data for an air preheat systems To cooler Diesel equipment were obtained,1 as well as typi285°F Top PA (Existing Kero/ Kero PA return cal cost factors for installations with either Desalter crude shell) return revamped or new fired heaters. Fig. 6 illus245°F Desalted crude trates the results for installing an air preheat AGO prod system to a new heater. The results show a Diesel prod significant reduced absorbed heater duty Diesel PA that is economically viable (20% IRR) as HVGO a result of cap-and-trade requirements. At AGO PA 475°F $40/metric ton, air preheat will be viable for Vac twr bottoms DSL a grassroots heater with an absorbed duty AGO PA just over 20 MMBtu/hr. With GHG valued Diesel Crude to at $15/metric ton, the threshold to meet the PA heater 20% IRR target is 40 MMBtu/hr.

Product 515°F 34.3 360°F 78.2 MMbtu/hr Steam Fig. 7 shows the economics for a revamp MMbtu/hr HVGO 2 (add 3 shells 48,000 4,000 ft of an existing heater with air preheat. The 13,500 ft2) lb/hr total install cost (TIC) factor is larger Diesel PA Quench for a revamp due to the amount of work return 490°F to VT required to modify the existing duct work, BFW VTB to stack and other potential changes. Larger HVGO PA coker heaters with an absorbed duty of just over Diesel AGO PA 150 MMBtu/hr absorbed duty would meet (Modified 27 MMBtu/hr 27 MMBtu/hr the 20% IRR target with GHG emissions AGO PA return service) 3,500 ft2 3,500 ft2 valued at $15/metric ton. A sensitivity 440°F analysis for natural gas prices at $6.50/ FIG. 5 Revamp crude preheat exchange network.

MMBtu (vs. $5.50/MMBtu for the base case) on an LHV basis showed benefits at HYDROCARBON PROCESSING JUNE 2010

I 37

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION

Air preheat economics–Grassroot

Air preheat economics–Revamp

60

45 40

$100

50

$100

35

$70

40

$70

30

$50

30

IRR, %

IRR, %

$30 $15

$50 $30

25

$15

20

$0/metric ton

20

$0/metric ton

15 10

10 5 0 20 FIG. 6

40

60 80 100 120 140 Process absorbed duty, MMBtu/hr

160

180

Grassroots heater air preheat economics with GHG emission costs included.

Motive steam

Vacuum tower

0 50 FIG. 7

100 150 200 Process absorbed duty, MMBtu/hr

250

Revamp heater air preheat economics with GHG emissions costs included.

Motive steam

Vacuum tower

Offgas

Cooling water

Offgas

Cooling water

To SWS Cooling water

To SWS

FIG. 8

Current vacuum steam ejector system.

FIG. 9

TABLE 1. Crude preheat exchanger network revamp results Heater

Base case

Revamp case

382.6

324.9

Steam in convection section, MMBtu/hr

36.1

28.3

Efficiency, %

87.0

87.4

481.25

404.1

0.0

–7.1

0

27,247

Absorbed duty, MMBtu/hr

Fired duty, MMBtu/hr Reductions Steam generated from VTB, MMBtu/hr GHG emissions, metric tpy

common seal drum. The motive fluid is 250 psig, 500°F mediumpressure (MP) steam and cooling water is available at 90°F. 38

I JUNE 2010 HYDROCARBON PROCESSING

Vacuum steam ejector system revamp with liquid ring compressor.

Depending on the trade-off between the costs of steam vs. electricity, it may be economical to replace an ejector stage with a liquid-ring vacuum pump (LRVP), referred to as a hybrid system.

The LRVP option has a lower steam use; thus lowering GHG emissions at the boilers.

The typical hybrid setup has an LRVP as the last stage of a multiple-stage system. An LRVP is not technically feasible as the first-stage. Replacement of the second-stage ejector with LRVP is not practical due to the high volumetric rate of the process gas. Table 2 summarizes the process data comparing the two cases.

The retrofit scope leaves the third-stage ejector and condenser abandoned in place, while the LRVP skid is located at grade. This would allow the third-stage ejector to be used when maintenance is performed on the LRVP. The seal-fluid configuration is closed-

PROCESS/PLANT OPTIMIZATION loop to minimize the seal water makeup rate. The typical skid includes seal fluid recirculation pumps, which were not needed by the notional refinery. Plot space is assumed to be available near the vacuum tower. An installed capital cost of $2.8 million estimate was developed. Fig. 9 shows the hybrid system.

The IRR for GHG emissions prices from $0/metric ton to $100/metric ton is shown in Fig. 10. The result shows that such a revamp would be marginal on fuel savings alone. However, with the additional cost that would result from GHG emissions in a proposed cap-and-trade program, the project has significantly improved economics.

product storage can be achieved via steam generation. Table 3 summarizes this case study process data.

A cost estimate was developed for the process revamp. In addition, 1,000 ft of new piping was included for each new feed line.

The estimated installed cost is $3.2 million; the revamp process schematic is shown in Fig. 12.

Results show that the economics with the assumed energy prices would make the project economically viable without GHG credits. Fig. 13 shows that the economics are enhanced significantly when GHG emissions costs are considered.

Power recovery. The FCC unit has a high-temperature flue gas stream from the regenerator that is typically used to make high-pressure (HP) steam. The flue gas is typically at 30 psig–35 psig with a temperature of 1,350°F. The flue gas is then let down

Diesel HDS hot feed with steam generation. The

notional refinery for this study has a diesel HDS unit that processes cold diesel feed from storage to produce ULSD. The cold feed is preheated in a low-pressure (LP) stripper bottoms exchanger. The reactor effluent heats the LP separator liquid (stripper feed) and generates steam prior to being collected in downstream flash drums. This steam generation is assumed to be economically justified during the conversion to ULSD. Fig. 11 is a schematic of this process.

For this study, generating additional 250-psig saturated MP steam was evaluated. A promising target for heat exchange was identified by utilizing a hot unit feed to reduce the duty on the unit feed/product exchanger. The additional cooling needed for

TABLE 3. Diesel HDS unit hot feed revamp results Generated

Base case

Revamp case

Steam, lb/hr

41,870

69,370

0

14,489

Reductions GHG emissions, metric tpy

TABLE 2. Vacuum steam ejector revamp results Steam, lb/hr Cooling water, gpm

Base case

Hybrid case

32,152

21,473

2,650

2,735

0

217

Electricity, bhp

IRR, %

Demand

Vacuum tower steam ejectors revamp

45 40 35 30 25 20 15 10 5 0 0

Reductions Boiler fired duty, MMBtu/hr

0

14.9

GHG emissions, metric tpy

0

6,649

SPECIALREPORT

FIG. 10

20

40 60 GHG price, $/metric ton

80

Vacuum tower steam ejector system revamp economics with GHG emissions cost included.

Wash water 120°F 250 psig stm 41,870 lb/hr

Recycle gas to scrubber and compressor 985 psig

Sour flash gas

126°F TC

100

LC

155 psig BFW

Treat gas

Sour water 1100

120°F

400°F

Sour offgas 40 psig

Quench gas TC

Cold diesel feed 45,000 bpsd

ULSD reactor

ULSD product 120°F FIG. 11

Quench gas Sour water

120°F

345°F

ULSD stripper

Sour wild naphtha product

Current diesel HDS system of notional refinery.

HYDROCARBON PROCESSING JUNE 2010

I 39

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION Wash water 120°F 250 psig stm 41,870 lb/hr

Recycle gas to scrubber and compressor 985 psig

Sour flash gas

126°F TC

LC

155 psig BFW

Sour water

Treat gas 1100

400°F FC

Tankage 120°F LCO

Quench gas

TC

FC

245°F Coker ULSD product

FC

230°F 45,000 bpsd

250 psig stm 28,370 lb/hr E-NEW ULSD product/steam generator 28-in./38-in. x 20 ft BKU 2,700 sq ft

315°F

ULSD stripper

Sour wild naphtha product

LC

BFW

Diesel HDS unit revamp with hot feed.

Diesel HDS hot feed revamp 120 100

IRR, %

Sour water

ULSD reactor

OSBL ISBL

300°F

120°F

80 60 40 20 0 0

20

40 60 GHG price, $/metric ton

80

100

Diesel HDS hot feed revamp economics with GHG emissions included.

Orifice chamber

33 psig 1,350°F

1.5 psig 1,350°F

Steam drum

Superheated steam 600 psig @ 600°F 163,000 lb/hr

FCC regenerator

Blowdown 8,000 lb/hr BFW 234°F 171,000 lb/hr To pollution control/stack 440°F

Flue gas cooler with economizer and steam superheat station 179 MMBtu/hr

FIG. 14

40 psig

TC

FC

Virgin diesel

FIG. 13

Sour offgas

Quench gas

300°F

FIG. 12

120°F

Existing FCC flue gas system.

to 1.5 psig before being routed to the flue-gas cooler and the emission controls (electrostatic precipitator or wet-gas scrubber). In the flue-gas cooler, 600-psig, 600°F HP steam is generated using a steam generator with an economizer and a superheat section.

Fig. 14 is the process schematic for this case.

A revamp of the FCC regenerator includes a power-recovery turbine and a third-stage separator to protect the turbine from catalyst fines. The hot flue gas is sent to the new power-recovery turbine bypassing the orifice chamber and work is extracted. The power-recovery turbine generates electrical power. The flue gas is cooled to 1,025°F at the outlet of the turbine and further cooled in the existing flue-gas cooler, which will generate less 600 psig, 600°F HP steam.

For our case study, the electric power produced by the turbine offsets power generated at a coal-fired plant. Depending on the type of projects identified to be eligible for offsets under the proposed Waxman-Markey bill, credits for the reduced GHG emissions could be provided to the refinery. Table 4 summarizes the process data for this case study.

An installed cost estimate of $40 million was developed for the revamp scope; it includes the power-recovery turbine, third-stage separator and other associated installation items. It is assumed that sufficient plot space would be available for this revamp. Fig. 15 is the process schematic for the revamp.

Limitations under the proposed Waxman-Markey bill that reduce GHG credits received from offsets are accounted for in the economics. The IRR for GHG emissions prices from $0/metric ton to $100/metric ton for this project is summarized in Fig. 16.

The results show that such a revamp would be marginal based on fuel savings alone. However, with the GHG valued at $20/metric ton, the threshold IRR of 20% would be achieved.

Summary. The presented case studies demonstrate a variety

of refinery energy conservation projects that could significantly 40

I JUNE 2010 HYDROCARBON PROCESSING

PROCESS/PLANT OPTIMIZATION

SPECIALREPORT

TABLE 4. FCC flue gas power recovery revamp results Generated

Base case

Steam, lb/hr

162,615

100,290

0

24,340

Electricity, bhp

33 psig 1,350°F

Reductions Boiler fired duty, MMBtu/hr

Orifice chamber Normally no flow used when turbine is down

Revamp case

0

1.5 psig 1,025°F

–21.8

GHG emissions reduced Direct, metric tpy

0

–20,679

Indirect, metric tpy

0

161,387

Total

0

140,708

improve economics when the costs of GHG emissions from a capand-trade program are considered. As a result of installing all of the projects presented here, a 9.5% reduction in GHG emissions from the notional refinery is achievable. Table 5 provides a net summary of the reductions in energy use and GHG emissions.

Observations from the various case studies show that care must be taken to evaluate all energy usage not only in the process unit of interest, but also changes in net refinery energy usage. Energy use in general will need a thorough review to ensure that all appropriate costs are included.

With the potential of offsets for power produced from coal, significant GHG reductions may be available from power recovery projects. Careful tracking of both electricity prices and available

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Power recovery turbine 24,000 SHP 80% adiabatic Eff Third-stage separator (protect turbine from Cat fines)

Steam drum Superheated steam 600 psig @ 600°F 100,000 lb/hr

Generator 18 MW

BFW 234°F 105,000 lb/hr To pollution control/stack 480°F

Flue gas cooler with economizer and steam superheat station 110 MMBtu/hr

FCC regenerator

FIG. 15

Blowdown 5,000 lb/hr

FCC flue gas power recovery revamp.

offset projects with the proposed legislation may provide some significant GHG reductions.

Various other opportunities are likely available at each refinery operation. An energy audit of the refinery operation is rec-

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I 41

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION

TABLE 5. Summary of energy and emission reductions from proposed energy conservation projects CO2 reduced, metric tpy

NG reduced, SCFH

Reduced boiler capacity, lb/hr

27,247

91,758

–6,720

6,649

16,376

10,679

Diesel HDS hot feed

14,489

35,714

27,500

0

No

3.2

FCC power recovery

140,708

–23,956

–62,325

18,158

Yes

40.0

Total

196,007

119,892

–30,866

17,996

Project Crude heater inlet temperature Vacuum tower ejectors

Power utility reduced, kWh

Offsets required

Capital cost, $ million

0

No

19.5

–162

No

2.8

65.5

IRR, %

If offsets are required, CO2 reduced includes the offsets awarded.

ACKNOWLEDGMENT This paper is revised and updated from an earlier presentation at the 2010 AIChE Spring National Meeting, March 21–25, 2010, San Antonio, Texas. In addition, the authors acknowledge the assistance provided by Graham Corp. for the analysis of the third-stage vacuum tower steam ejectors with a liquid-ring compressor; budgetary costs for air preheat equipment provided by Tulsa Heaters, Inc., and OnQuest; and economic factors for revamp and grassroots air preheat economics provided by Furnace Improvements.

FCC power recovery revamp

45 40 35 30 25 20 15 10 5 0 0

FIG. 16

20

40 60 GHG price, $/metric ton

80

100

FCC flue gas power recovery revamp economics with GHG emission costs included.

ommended so that the options presented, as well as additional opportunities, can be identified. The economic analysis should include the cost of GHG emissions since previously uneconomic projects may become viable as a result of proposed cap-and-trade for GHG emissions. HP

1

LITERATURE CITED US Energy Information Administration (EIA), “Emissions of Greenhouse Gases in the United States 2008,” Department of Energy, December 2009.

Ian M. Glasgow is a senior technical professional for the downstream process engineering group for Mustang. He is responsible for basic and detailed process design. Prior to Mustang, he was a research associate in process integration in the pulp and paper industry and research chair at Ecole Polytechnique de Montreal. He holds a BS degree in chemical engineering from the University of Cincinnati and an MS degree in chemical engineering from Auburn University.

Stan Polcar is a process manager in the downstream process engineering group for Mustang. He is responsible for leading process design teams on work ranging from front-end studies to detailed engineering projects. Mr. Polcar has been with Mustang for 13 years. Prior to Mustang, he was with Litwin Engineers and Constructors and with Pritchard. He holds a BS degree in chemical engineering from Case Western Reserve University.

Earl Davis is a principal technical professional in the downstream process engineering group for Mustang. He is responsible for conceptual and detailed process design for refinery projects. Prior to joining Mustang, he held similar positions with Litwin Engineers and Constructors and with Davy McKee Corp. He holds a BS degree in chemical engineering from Youngstown State University.

Tram Nguyen is a senior consulting technical professional for the downstream

GJP4.5e10

process engineering group for Mustang. She is responsible for basic and detailed process design. She holds a BS degree in chemical engineering from Rice University.

Jerry Price is a process manager in the downstream process engineering group

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for Mustang. He is responsible for leading process design work ranging from frontend studies to detailed engineering projects. Mr. Price has been with Mustang for 11 years. Prior to Mustang, he was with Litwin Engineers and Constructors, Parsons, Badger and Foster Wheeler. He holds BS and MS degrees in chemical engineering from the University of Houston.

Chris Stuecheli is a senior consulting technical professional for the downstream process engineering group for Mustang. He is responsible for basic and detailed process design. Mr. Stuecheli has been with Mustang for 8 years. Prior to Mustang, he was with Litwin Engineers and Constructors and with Technip. He holds a BS degree in chemical engineering from the University of Texas at Austin.

R.E. (Ed) Palmer is the manager of downstream process engineering for Mustang. He is responsible for process design and marketing support for all refining, petrochemical and chemical projects. At Mustang, he has led numerous studies, technology evaluations, and projects relating to clean fuels production. He has authored numerous articles and industry meeting presentations relating to petroleum refining. Prior to assuming his current position, he spent 23 years with Litwin Engineers and Constructors in a variety of assignments. Mr. Palmer was employed as a refinery process engineer for Conoco in Oklahoma. He holds a BS degree in chemical engineering from the University of Missouri, Rolla.

PROCESS/PLANT OPTIMIZATION

SPECIALREPORT

Low-cost advanced process control project captures energy savings in utilities area This low-cost project resulted in over $300,000 per year in benefits E. CHANG and M. VIDUCIC, LyondellBasell Industries, Corpus Christi, Texas

T

he rise in energy prices presents new challenges to the manufacturing facility profitability. The hydrocarbon and chemical processing industry is particularly affected due to the large amount of energy required to operate these facilities. While advanced process control (APC) has been widely applied over the past several decades to the process units of these industries, the utilities area has received relatively little attention due to the lower potential for benefits. Today’s high energy prices, and the pressure to reduce emissions, are making APC projects in the utilities area much more attractive and necessary to stay competitive.

At LyondellBasell’s Corpus Christi Olefins Complex in Corpus Christi, Texas, an APC project was recently completed to optimize the dilution-steam header pressure. The project cost and implementation time were minimal, and the success and operator acceptance of the new controller were exceptional. The realized benefit is in excess of $300,000 per year.

Impact of high energy costs. Whether it is fuel gas to fire

a furnace, electricity to run a pump or steam to drive a turbine, the second highest manufacturing cost in a petrochemical facility is energy. While the energy cost is an uncontrollable variable, energy usage can be monitored in great detail and controlled to an optimum.

In the past few years, the price of natural gas has increased 300%, topping out in 2008. A downward trend has occurred since this peak. The tremendous fluctuation in energy price demonstrates the dramatic impact that energy has on petrochemical plant profitability. After feedstock, energy is the highest manufacturing cost of an olefins complex and can account for approximately 40% of the operating cost.

Understanding supply and demand of all energy sources is critical in managing the impact that energy costs have on a plant. Detailed steam and fuel balances are important monitoring tools. Energy usage can be monitored by developing a daily energy-tracking sheet that accounts for all energy usage. This can be looked at from an overall view or can be drilled down to individual equipment. Furnace thermal efficiencies, steam usage by major turbines or overall fuel gas and electricity usage in the complex can be captured. With a good energy usage view, variances can be monitored and optimized. Cost reductions can

be achieved through various means, ranging from major capital expenditures to zero-cost improvements.

Examples of capital expenditures include upgrading or retrofitting major steam turbines or fired heaters to gain valuable efficiency. Examples of zero-cost improvements include reducing excess reflux on distillation towers or reducing excess oxygen in fired heaters. The age of a petrochemical complex can determine how difficult it is to find energy-reduction opportunities and how costly they might be to implement. Typically, the older the complex, the easier it is to find opportunities. However, these opportunities tend to cost more due to obsolete technology. Newer sites will tend to have newer technology, making finding opportunities a bit more difficult. In either case, there is almost always some fairly easy, low- to zero-cost options to be found. Optimizing the dilution-steam header pressure is one such example.

The dilution-steam system. The process flow of the dilution-steam system is shown in Fig. 1. Dilution steam is superheated steam that serves several purposes in a cracking furnace.

It is injected into the furnace feed stream prior to the second Dilution steam

Hydrocarbon feed

Oily water separator Cracking furnace Primary fractionator FIG. 1

Low-press stripper

High-press stripper

Dilution-steam system.

HYDROCARBON PROCESSING JUNE 2010

I 43

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION

hydrocarbon pass, vaporizing and diluting the hydrocarbon feed and taking it through the dry point very quickly. This helps prevent convection-section tube coking and fouling. Dilution steam lowers the furnace feed partial pressure so that more effective thermal cracking is obtained. It prevents long-chain hydrocarbons from rejoining in the cracking process, which would lead to forming undesirable compounds in the furnace effluent.

To cracking furnaces FI

PIC

Feed from LP stripper

TIC FI

FI

200-psig 200-psig steam steam

FI

Prim. fract. quench oil

200-psig steam

Zero-cost energy savings opportunity. The dilutionsteam system optimization “problem” is fairly straightforward— reduce the header pressure as much as possible (thereby saving intermediate pressure steam) while ensuring that the cracking furnaces have sufficient steam to operate. The individual furnace dilution-steam flow controller outputs are the constraints.

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High-pressure water stripper.

Visual Level Indicators

FIG. 2

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I JUNE 2010 HYDROCARBON PROCESSING

The furnace effluent, containing cracked hydrocarbons and steam, flows to the primary fractionator, where the hydrocarbons are separated and steam is condensed. A side-draw of condensed dilution steam and hydrocarbons is routed to a liquid separator, referred to as the oily water separator. In the oily water separator, the oil and water are allowed to settle and separate. A portion of the water is removed from the separator, stripped of any remaining hydrocarbons (by the low-pressure water stripper) and vaporized (by the high-pressure water stripper) to generate steam, which is then routed back to the furnaces for use as dilution steam.

Fig. 2 shows the control system for the high-pressure stripper.

Heat to generate steam comes from two sources: the primary fractionator bottoms quench-oil circulation loop and 200-psig steam. The dilution-steam header pressure is controlled using a split-range controller. Normally, header pressure is maintained by regulating the 200-psig steam rate. If there is insufficient reboiler duty to meet dilution-steam demand, live 200-psig steam can be injected directly into the header to achieve the desired header pressure. A temperature controller on the highpressure stripper overhead injects live 200-psig superheated steam to ensure that the overhead steam is superheated.

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PROCESS/PLANT OPTIMIZATION To capture this energy savings, two zero-cost options were considered. The first was to simply ask that operations keep a closer eye on the dilution-steam header pressure and maintain the pressure as low as possible. The second, more aggressive, option was to somehow automate this pressure minimization.

Operations was hesitant about the automation option because doing so might cause instability for all the furnaces on the header. It was agreed that pressure minimization would be done manually by the operators. To aid the operators, an online calculation was configured in the distributed control system (DCS) to estimate which furnace was most limiting and how

FIG. 3

Multivariable controller independent variables display.

SPECIALREPORT

much pressure reduction was possible (how much minimization “room” was available).

Manual pressure minimization was tried for one year and its success was limited. Operations made adjustments to the header pressure, but these adjustments were infrequent since they had many other activities of higher priority occupying their time. Also, the “room” available was still quite large and so much of the savings were left uncaptured. So it was obvious that automating this optimization problem was the only effective method by which to capture this energy savings. Operations agreed to give it a try.

FIG. 4

Multivariable controller dependent variables display.

Select 160 at www.HydrocarbonProcessing.com/RS

HYDROCARBON PROCESSING JUNE 2010

I 45

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION

Automation to increase savings. In January of 2009, a small multivariable controller (MVC) was commissioned to optimize the dilution-steam header pressure. The objective of this controller was very simple—reduce the header pressure until the most limiting furnace’s dilution-steam flow controller output reached its maximum limit, and increase the header pressure when necessary to keep this controller within control range.

Although an MVC was utilized for this application, this constraint problem could probably have been solved just as

FIG. 5

Three-day MVC performance trend.

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easily with a series of DCS override controllers but MVC was chosen in part because the necessary models were already available as part of the cracking furnace MVCs. Therefore, no step testing was required for this project. (In the cracking furnace MVCs, the dilution-steam header pressure is a disturbance variable to the dilution-steam flow controller output constraint included in each furnace controller.)

The dilution-steam MVC is a small controller, with only one manipulated variable (MV) and 17 controlled variables (CVs). No disturbance variables were included. The single MV is the dilution-steam header pressure controller setpoint. The CVs are the individual dilution-steam flow controller outputs.

Of the 15 cracking furnaces, 13 have single-dilution-steam flow controllers and two have dual flow controllers (due to dual feed passes). MVC execution is once per minute. Figs. 3 and 4 show the independent variable and dependent variable DCS screens, respectively.

Logic was included to automatically exclude a furnace from the controller if the furnace was not online, or if the furnace’s dilution-steam flow controller was in manual mode. This is illustrated in Fig. 4, where the controlled variables for furnaces C, D and E are turned off because these furnaces are not online. Furnace C is completely down and furnaces D and E are on hot steam stand-by.

The controller was tuned to move the header pressure very slowly so that pressure moves to minimize the pressure or to relieve a constraint did not adversely affect the furnaces. To make it easier for operators and engineers to determine which furnace was limiting, logic was included to indicate the most limiting furnace based on the dependent variable’s steady-state constraint status.

Consideration was given to adding each furnace’s dilutionsteam flow controller setpoint as disturbance variables. In the end, it was determined that this would be of little value toward improving controller performance given that the dilution-steam flow for each furnace is manipulated by the furnace’s MVC, and the moves are typically very small and gradual. So any constraint violations resulting from dilution-steam setpoint changes should be handled adequately by the dilution-steam header MVC on error feedback.

The work to develop and commission the controller was all performed by in-house personnel. The infrastructure was already in place, so there was no cost incurred for instrumentation, MVC software license or computer hardware. Aside from internal manpower resources, the project was completed at no cost and took only one week to complete.

Controller performance. Fig. 5 is a 3-day plot showing the dilution-steam flow controller output of one furnace (purple trend lines) and the dilution-steam pressure controller setpoint (blue trend lines). The heavy purple line indicates the CV’s upper limit, set at 85% output. The two heavy blue lines indicate the MV upper and lower limits. Initially, the MVC is able to gradually minimize the header pressure. As it does over the first day, the dilution-steam flow controller output gradually rises until it hits its upper limit. Upon hitting the upper limit, the MVC stops lowering the header pressure and over the next two days the pressure rides on this furnace’s flow controller output upper limit since this was the most limiting furnace.

The MVC onstream time since commissioning (defined as the percent of time that the controller is in service), has been

PROCESS/PLANT OPTIMIZATION excellent at just over 99%. Initially, the dilution-steam flow controller output upper limits were set at 75% output. Once the controller was proven and operations felt comfortable with its performance, this upper limit was raised to 85% and the header pressure lower limit was lowered from 98 psig to 97 psig. An upper limit of 85% output was chosen (and not higher) to ensure that in a constrained situation, there would be some room to increase the dilution-steam rate if necessary, while the header pressure MVC adjusted the header pressure slowly. Since the MVC was placed in service, the header pressure has been riding an output constraint most of the time and has rarely been sitting at its lower or upper limits. In effect, the MVC has been optimizing and capturing benefits most of the time—something that was not possible manually.

Benefits achieved. The dilutionsteam header pressure (prior to MVC) was historically maintained at approximately 110 psig. This was a convenient setting for operations since it ensured sufficient dilution steam for all furnaces at varying hydrocarbon feed rates. However, the price for this pressure setting was generating excess dilution steam. Implementing an MVC automated header pressure optimization and matched steam make with furnace demand. Since MVC implementation in January 2009, the dilution-steam header pressure average has been reduced to 102 psig. This essentially no-cost solution resulted in an 8-pound pressure reduction, netting a savings of approximately $300,000 per year (73,000 mm Btu/year).

SPECIALREPORT

Inc., and Aspen Technology, Inc. He has 20 years of advanced process control and real-time optimization experience, and has commissioned controls at refineries, petrochemical and chemical facilities in the U.S. and overseas. Mr. Chang holds a BE degree in chemical engineering and computer science from Vanderbilt University, and is a registered professional engineer in Texas.

Mark Viducic is a senior process engineer for LyondellBasell Industries, in Corpus Christi, Texas. Prior to joining LyondellBasell Industries, he worked for Formosa Plastics Corporation.

Mr. Viducic has 11 years of olefins process experience and has spent the last few years focusing on energy-usage reduction. Mr. Viducic holds a BS Degree in Chemical Engineering from Texas A&M University, College Station.

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Final considerations. An understanding of energy usage and recognizing an opportunity to reduce energy consumption led to this very successful project.

This same approach could be applied to other steam headers. While many header systems include a large number of users, this does not necessarily mean a large and complex control application. Typically, only a handful of users or constraints limit how much a header pressure can be optimized. Finding the most-limiting constraints is the key to a successful control scheme. Also, consideration must be given to how aggressively the control scheme moves the header pressure. Automatic adjustments must not adversely affect the control performance of the individual steam users. HP Ed Chang is a principal process control engineer for LyondellBasell Industries in Corpus Christi, Texas. Prior to joining LyondellBasell, Mr. Chang worked for Celanese AG, Setpoint, Select 162 at www.HydrocarbonProcessing.com/RS 47

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H58646

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PROCESS/PLANT OPTIMIZATION

SPECIALREPORT

Modify your vacuum-tower transfer line to increase benefits Designing the vacuum unit is a challenge. Revamping the transfer line can make this unit more profitable.

R. YAHYAABADI, Esfahan Oil Refining Co., Esfahan, Iran

T

ougher economic conditions challenge refiners to be profitable under severe situations. In refineries, vacuum units greatly impact profitability. For the vacuum unit, profitability largely depends on product yields and unit run length.

Many published articles provide valuable guidelines on how to operate to obtain higher yields and longer run lengths. These articles mostly discussed the vacuum tower critical sections.1 However, there are economic opportunities in debottlenecking the vacuum transfer line. The transfer line is one of the important critical sections on a vacuum unit. We will discuss the effects of the transfer line on the vacuum tower operating conditions and products yields and specifications, as well as modifications to achieve more yields.

and HVGO yields, those product specifications limit VGO applications. In other words, increasing the VGO yield within the specifications for catalytic processing is the goal LVGO Vacuum for all vacuum units. But high column VGO yield within the specifications of catalytic unit feeds requires good engineering for HVGO the vacuum unit.

Conversely, to increase VGO Wash oil Wash zone yields, severe operating conditions are necessary for the vacCollector tray uum unit. Such conditions can Vapor horn limit the unit run length. For Slop wax Flash zone Transfer line example, vacuum units with high Fuel VGO yields (severe conditions) and long run lengths (operatSteam ing four to five year continuous runs) require perfect design.

VRES While higher VGO yields mean higher TBP cutpoint, severe operating conditions also Typical vacuum unit at a refinery. involve higher temperatures and lower pressures. For vacuum profitability of the vacuum unit is mainly processing, high temperature related to LVGO and HVGO yields and requires special attention to avoid coking their specifications. within the system.

VGOs are typically sent to catalytic units Processing at lower pressures while the for conversion into higher value products. system is working at higher temperatures is Consequently, there are restricted limita- difficult; higher temperatures involve hantions over the VGOs sent to the catalytic dling higher gas volumes that will be manunits. Common specifications for VGOs aged by the vacuum-producing system. Due include: metal content, micro carbon to these operating conditions and criteria, residue (MCR) and asphaltenes; all limit designing a vacuum unit while considering VGO usage in the catalytic units. Process economic aspects is a challenge. and equipment designs that minimize distillation tails will reduce metals content.3 Critical points in the vacuum unit.

Minimizing HVGO metals can dramati- In the vacuum units, some critical points cally increase catalyst service life.4 act as determinative factors. All of these While economic aspects force operat- points must be considered during design, ing the vacuum unit to increase LVGO especially for higher VGO yields. The most To Vacuum system

Feed

Fired heater

FIG. 1

Vacuum units. A refinery

vacuum unit has several major sections, including: vacuum heater, transfer line, vacuum tower and vacuum producing system. Fig. 1 shows a typical refinery vacuum unit. There are different vacuum tower types; detailed discussions for each design have been presented in earlier articles. 2 Among the different types of vacuum towers, the configuration (shown in Fig. 1) is the most common.

In this arrangement, feed from the atmospheric tower bottom is distilled into two valuable vacuum gasoil products— light vacuum gasoil (LVGO) and heavy vacuum gasoil (HVGO)—and the bottoms known as vacuum residual (VRES). The

HYDROCARBON PROCESSING JUNE 2010

I 49

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION

common problem is coke formation in the fired heater and wash zone.1 Coking in the heater outlet is a common.5 Coke forms inside the radiant section tubes of the vacuum heater as the oil film flowing along the inside of the tube exceeds the temperature and residence time needed to initiate thermal cracking.5 So, controlling the oil-film temperature and residence time is very essential when minimizing coke make.5 The wash zone can contribute to coke

make due to poor HVGO product quality, low HVGO yield and unscheduled outages to replace packing.6 Inadequate washzone liquid flowrate is a primary cause for coking.7 To mitigate coke make, sufficient wash oil flow is needed to keep the middle of the packed bed wetted. Otherwise, high residence-time stagnant zones are created.8 Coke forms in the middle because it is the only part of the bed that is not wetted.8 Coking in the middle of the wash zone has been discussed in other articles.1

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Vacuum-unit product yields and critical operating conditions must be accurately predicted.8 System features include: heater outlet, transfer line, flash zone, collector tray below the wash section and wash section column internals.8 The transfer line, flash zone and wash-section designs do influence the coking rate in the wash section internals.9 Effect of vacuum tower transfer line.

Vapor and liquid feed enters the column at velocities as high as 380 ft/sec–400 ft/ sec.6, 8, 10 The vapor phase contains small droplets of VRES that are generated in the transfer line. The droplet size is too small to allow settling in the transfer line due to the very high velocity of the vapor.6, 8, 10 Accordingly, the flash zone and wash section must remove the entrainment.6 The flash-zone vapor horn and flash zone help remove larger droplets and distribute the rising vapor across the cross-section of the column.6 By uniformly distributing the vapor, high velocity areas are minimized, thus allowing the packing to remove essentially all small residue droplets.6 In the vacuum unit, the transfer line, flash-zone vapor horn and wash-section internals determine the level of entrainment.3 The quantity of entrainment for a unit varies according to the flash-zone design and height, transfer-line velocity, etc.11 A poorly designed transfer line with a high-pressure drop at the column inlet nozzle can generate a fine mist that is difficult to remove.3 Yet, entrainment can be almost eliminated through prudent transfer line and column internal designs.3 Detailed studies on the entrainment from the flash zone to the wash zone have been presented elsewhere.1 While entrainment from the flash zone contains high metals, concarbon and asphaltenes, the amount of entrainment should be minimized as much as possible.1 Transfer line, flash zone and washsection designs influence the HVGO concarbon, metals and asphaltenes content.9 Estimating the heater and transfer-line pressure profile accurately requires a model capable of rigorous tube-by-tube heat transfer and accurate two-phase flow calculations.8 Calculated phase regimes in the transfer line are either stratified or stratified wavy, as shown in Fig. 2.9,10 Two phase system. Stratified phases cause the liquid and vapor to have poor mass and energy exchange across the interface. 8, 10 Thus, liquid-vapor contact is poor.10 Since the transfer line consists of large diameter piping, the liquid and vapor phases separate in the horizontal section;

PROCESS/PLANT OPTIMIZATION 170

Rate, m3/hr

150

Vapor

Liquid

A transfer line with phase separation.

FIG. 2

Flash

Transfer line liquid

Transfer line

Non-ideal stage for FZ

Splitter Steam

Distilltion point,°C

Entrainment

Furnace outlet Flash

2.50 Wash oil rate O/F rate Dryout ratio

130 110

70

2.42

50

2.40 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 Degree of superheating, °C Effect of the transfer line non-idealities on several operating conditions with entrainment.

600 590 580 570 560 550 540 530 520

HVGO EP HVGO 95% VRES 5%

2

Non-equilibrium in the transfer line with a non-ideal stage for the flash zone and entrainment in the wash zone.

the vapor flows along the top of the pipe, and the liquid flows across the bottom.8, 10 The transfer-line vapor becomes superheated due to pressure reduction as the two phases approach the flash zone.8 Phase separation causes the superheated vapor to flow through the top of the pipe and colder liquid to be on the bottom.9 The vapor and liquid entering the flash zone are not in equilibrium.8, 10 Assuming that the liquid and vapor entering a vacuum column flash zone are in equilibrium is a critical mistake.8 The transfer-line phase separation increases the amount of wash oil needed to prevent coking; the wash oil vaporizes more of the wash liquid.8 In reality, accounting for transfer-line phase separation raises the wash oil flowrate by 200%–300% over conventional modeling practices assuming that the liquid and vapor leaving the transfer line is in equilibrium.10 Evaluating different cases. As mentioned before, the sections that are important and critical and that must be accurately simulated are the heater outlet, transfer line, flash zone and wash zone. Other parts of the vacuum column are straightforward and well understood. The whole unit can be simulated. But analyzing and evaluating

2.46 2.44

3

4

5

VRES FIG. 3

2.48

90

FIG. 4

Wash oil Transfer line vapor

2.52 Dryout ratio, vol.

Vapor/liquid phase separation

SPECIALREPORT

FIG. 5

6

7 8 9 10 11 12 13 14 15 16 17 18 Degree of superheating, °C

Effect of the transfer line non-idealities on several product specifications with entrainment.

different models, only particular sections of the vacuum unit will be considered. To evaluate several different cases, simulation models were done following these rules: • Two theoretical stages were applied for the wash bed • Heater outlet temperature was set for a TBP cutpoint of 1,000°F (538°C) on HVGO cut. It was found that the heater outlet was within the normal range for such a TBP cutpoint.

• All slop wax was sent to the top of the stripping section • Flash-zone pressure, transfer-line pressure drop and, consequently, heater outlet pressure were fixed for all cases • Non-equilibrium flash zone was applied in all cases1 • Amount of entrainment from the flash zone is the same in all cases • Entrainment from the flash zone comes up to the middle of wash bed in all cases1 • Tower top pressure and temperature for all cases are the same • Same amount of stripping steam was used in all cases • Same number of theoretical stages was assumed for the stripping section • Minimum wetting rate of 0.15 gpm/ ft2 for the wash zone was set for all cases

• Multiple unit operations were used to model the non-ideality in transfer line (non-equilibrium system).

As discussed earlier in detail, considering all of the above items and non-idealities provide a model as shown in Fig. 3.1 Now, with this model, the effects of non-idealities in the transfer line on some important operating parameters and HVGO yield and specifications could be accurately studied and evaluated. As non-idealities in the transfer line produces superheated vapor, the amount of nonidealities could be shown as the degree of superheating of the produced vapor.

The amount of vapor superheating shows how much the transfer line is distant from the equilibrium or ideal transferline conditions. For the presented model, some operating conditions and product specifications have been calculated and are shown in Figs. 4 and 5.

With increasing non-idealities in the transfer line, the superheated vapor vaporizes some wash oil. With constant heater duty and product yields, the amount of wash oil and over flash (O/F) decreases as shown in Fig. 4. Note: The dry out ratio (wash-oil rate to over flash rate) increased.

Fig. 5 shows that while the transfer line is distant from the equilibrium transfer HYDROCARBON PROCESSING JUNE 2010

I 51

2

FIG. 6

3

38

640 620 600 580 560 540 520 500 480

33 28 23

2

3

4

5

HVGO 95%

Effect of the transfer line non-idealities on several operating conditions with no entrainment.

FIG. 7

Mixing device

Fired heater

FIG. 10

FIG. 8

FIG. 9

Vapor and liquid mixing device at the end of the transfer line.

Installing vanes as a mixing device at the end of the transfer line.

line conditions, the HVGO 95% and end point (EP) increase and the VRES 5% decreases. The non-idealities in the transfer line cause some losses in fractionation. As mentioned before, increasing the HVGO 95% and EP are not favored due to processing limits for the downstream catalytic units. A reduction of 5% in the VRES is not desirable when product is sent to an asphalt plant.

52

I JUNE 2010 HYDROCARBON PROCESSING

18 7 8 9 10 11 12 13 14 15 16 17 Degree of superheating, °C

HVGO EP

VRES 5%

HVGO distillation tail

Effect of the transfer line non-idealities on several product specifications with no entrainment.

Feed

The place for vapor and liquid mixing device installation

6

HVGO distillation tail, °C

120 105 90 75 Wash oil rate O/F rate 60 Dryout ratio 45 30 15 0 4 5 6 7 8 9 10 11 12 13 14 15 16 17 Degree of superheating, °C

Distillation point, °C

160 140 120 100 80 60 40 20 0

PROCESS/PLANT OPTIMIZATION

Dryout ratio, vol.

Rate, m3/hr

SPECIALREPORT

Using two transfer line and mixing devices near the tower inlet.

Entrainment issues. The effects of entrainment on vacuum-tower performance have been deeply discussed elsewhere.1 Entrainment from the flash zone is not favored in vacuum towers; it can cause several critical problems and difficulties on operating conditions and product specifications. Entrainment can plug the wash section due to coke particles generated by thermal cracking.1 Plugging the wash section causes low quality and yields of VGOs and, consequently, reduces plant profitability.1 Plugging the wash section is one of the worst events in a vacuum unit; it requires a shutdown to replace the bed packing.1 Entrainment from the flash zone is undesirable in vacuum towers and should be minimized to the lowest levels possible.

Due to entrainment issues, many designs and methods have been proposed for the vacuum-tower flash zone to eliminate entrainment. The effects of transfer line non-idealities on the operating conditions and products specifications for a vacuum tower that equipped with a no entrainment flash zone are shown in Figs. 6 and 7.

Although the treatment of the curves in Fig. 6 are similar to Fig. 4, changes in the curves for a no-entrainment flash zone are higher. For a no-entrainment flash zone, the wash-oil rate and over-flash rate sharply decrease when the non-idealities of the

transfer line increase. The sharp increase (or jump) of the dryout ratio after 12°C superheating is interesting. Comparing Fig. 6 with Fig. 4, we find that the wash oil and over flash of the tower with no entrainment are more sensitive to non-idealities in the transfer line than for the tower with entrainment.

Regarding the HVGO 95%, EP and VRES 5% specifications, the same treatments occurred when comparing Fig. 7 to Fig. 5. Increasing the HVGO 95% and the EP and decreasing the VRES 5% with an increase in superheating are sharper than the prior case. In this case, with increased superheating above 12°C, the HVGO EP sharply increases so that the HVGO distillation tail is raised. In the tower with the no entrainment flash zone, the HVGO 95%, EP and VRES 5% are more sensitive to superheating than for the case with entrainment.

Approach to equilibrium in the transfer line. Regarding the transfer

line, any attempt to achieve higher profits (higher yields or better HVGO quality) should be directed at achieving equilibrium in the transfer line as much as possible. An equilibrium transfer line means that the produced vapor and liquid are in equilibrium at the outlet.

As mentioned earlier, the liquid and vapor separate in the horizontal section of the transfer line. To achieve an equilibrium transfer line, the vapor and liquid should be mixed together as much as possible before exiting. Any attempt to approach equilibrium in the transfer line may be summarized by mixing the vapor and liquid at the end of transfer line, as shown in Fig. 8. Since the transfer line outlet velocity is too high, the mixing device should have special specifications, such as: • Resistance to erosion corrosion • Low pressure drop

PROCESS/PLANT OPTIMIZATION 6

• Special attention to avoid mist production.

For a typical mixing device, installing vanes at the end of the transfer line can be considered, as shown in Fig. 9.

The number of vanes and type determine the amount of mixing. CFD analysis can provide a good understanding for such arrangements. The important problems in this configuration are: • Mechanical strength of the vanes due to high velocity of the phases • Mist production.

In some cases, these problems may not be resolved, and other arrangements are considered. In some cases, two parallel transfer lines may be used, and the two transfer line streams will be mixed near the tower, as shown Fig. 10. A venturi-type mixing device can be used in this method.

As the vacuum transfer lines already consist of too large diameter pipes, this method may require higher investment.

For a revamp case, the method shown in Fig. 9 may be considered. From a design point of view, the arrangement shown in Fig. 10 may be the optimum choice. Using a flash drum near the tower inlet may be an option. But due to the high vapor volume in the vacuum condition, a very large flash drum will be required.

Golden, S. W., “Revamps: maximum asset utilisation,” Petroleum Technology Quarterly, Q4 2005, p. 37.

7 Golden, S. W., “Troubleshooting vacuum unit revamps,” Petroleum Technology Quarterly, Q2 1998, p. 107.

8 Barletta, T., and S. W. Golden, “Deepcut vacuum unit design,” Petroleum Technology Quarterly, Q4 2005, p. 91.

9 Hanson, D. and M. Martine, “Low capital revamp increases vacuum gas oil yield,” Oil & Gas Journal, March 18, 2002.

10 Martin, G. R., “Vacuum unit desig effect on operating variables,” Petroleum Technology Quarterly, Q2 2002, p. 85.

11

1

2

3

4

5

Golden, S. W., N. P. Lieberman and E. T. Lieberman, “Troubleshoot vacuum columns with low-capital methods,” Hydrocarbon Processing, July 1993, p. 81.

Reza Yahyaabadi is a senior process engineer for Esfahan Oil Refining Co. (EORC), Esfahan, Iran.

He has 21 years of experience in process engineering, process revamps, debottlenecking and simulation, and has a BS degree in chemical engineering from the Esfahan University of Technology.

More options. Effects of vacuum tower transfer line performance on HVGO 95%, EP and VRES 5% have been discussed.

Also, several arrangements to an ideal transfer line or equilibrium transfer line have been investigated. The equilibrium transfer line is a complex issue. Using CFD models can help determine the best final design. Another important parameter is the transfer-line pressure drop when using any arrangement to reach the equilibrium transfer line. Based on the selected method, the transfer-line pressure drop and its effect on the fired heater should be calculated. HP LITERATURE CITED Yahyaabadi, R., “Consider practical conditions for vacuum unit modeling,” Hydrocarbon Processing, March 2009, pp. 69–74.

Yahyaabadi, R., “Improve design strategies for refinery vacuum tower,” Hydrocarbon Processing, December 2007, pp. 106–112.

Golden, S. W., T. Barletta, and S. White, “Vacuum unit design for high metals crudes,” Petroleum Technology Quarterly, Q4 2007, p. 31.

Golden, S., “Canadian crude processing challenges,” Petroleum Technology Quarterly, Q4 2008, p. 53.

Golden, S. W., and T. Barletta, “Designing vacuum units,” Petroleum Technology Quarterly, Q2 2006, p. 105.

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Advanced process control: quick and easy energy savings Here are some of the energy savings benefits routinely produced by implementing advanced process control P. KESSELER, Invensys Operations Management, Houston, Texas

gas: $13/MMBtu; steam: $15/Mlb) and a more average value for the early 2000s (natural gas: $8/MMBtu; steam: $10/Mlb).

Mechanics of APC energy savings. In a nutshell, APC produces benefits by pushing constraints and performing more consistently and accurately than an operator can acting alone. A very common objective of APC applications is to lower the unit average operating pressure, which will often yield lower energy 16 14 Dollars per tcf

I

n today’s manufacturing environment, there is an urgency to increase operating efficiencies, and to do it quickly. One area of improvement that can produce immediate results is reducing energy consumption. It’s good for the environment and it’s good for the bottom line. “Energy management,” therefore, has become a common best practice, but there is more there than meets the eye. Typically, it implies rigorously modeling all or a major portion of the plant, coupled with the use of real-time optimization technology. While this approach has been used successfully, there are other simpler, faster options for reducing energy consumption in a manufacturing plant.

For example, advanced process control (APC) technology, which has been used to increase unit capacity or yield of the plant’s most valuable products, has also been successfully applied to produce energy savings in manufacturing facilities for the past two decades or more. This article will present some of the energy savings benefits routinely produced by implementing APC and show how applying these best practices can be quite lucrative at today’s energy prices.

12 10 8 6 4 1975

1980

1985

1990

1995

2000

2005

2006

Source: US Energy Information Administration

FIG. 1

Monthly US price of natural gas sold to commercial consumers.

14 12 Close 10 8 6 4

$3.945 03/04/2009

02/02/2009

12/31/2008

12/01/2008

10/30/2008

10/01/2008

09/02/2008

08/01/2008

07/02/2008

06/03/2008

05/02/2008

04/03/2008

03/04/2008

2 02/01/2008

was really beginning to hit its stride in the late ’80s and early ’90s, the financial benefits of energy savings were typically small when compared to unit capacity or product yield increases. For instance, the total US dollar benefit for an APC project applied to an oilrefining unit was typically in the $800,000 to $1,500,000 per year range. Of that total, the energy savings averaged $100,000 to $200,000. In contrast, recent APC projects have yielded savings of $400,000 to $900,000 in reboiler steam energy when applied to just a couple of refinery distillation towers.

The technology hasn’t changed much over the years, but the cost of energy has. A review of United States Department of Energy documents (Fig. 1) shows that the price of natural gas has increased from about $4.5/Mscf in 1988 to about $13.0/Mscf in 2006. If this cost was applied to a boiler with 80% efficiency, the cost of 150-psig steam has increased from approximately $6/ Klb of steam to $17.5/Klb of steam. Similarly, the refinery fuel product price has approximately quadrupled from 1991 until 2007. These increases account for most of the observed increase in energy-related benefits from APC applications. More recently, energy prices have tumbled back to the levels of 10 years ago (Fig. 2). For the purposes of this article, the dollar value of energy savings will be presented both in terms of peak energy prices (natural

$/MMBtu

Historical APC energy reduction benefits. When APC

Feb. 1, 2008 – Mar. 6, 2009 FIG. 2

NYMEX natural gas futures close (front month).

HYDROCARBON PROCESSING JUNE 2010

I 55

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION

requirements. APC applications also commonly reduce operator conservatism, some examples of which include: stack flue-gas O2 content higher than necessary for safe operation and enviTABLE 1. Key gas plant operating data Before APC After APC Aft. – Bef. Delta Debutanizer reboiler steam (Mlb/hr)

20.3

14.4

–5.9

Naphtha splitter pressure (psig)

31.4

24.7

–6.7

Naphtha splitter reboiler steam (Mlb/hr)

13.2

10.0

–3.3

10/21/2003 11/18/2003 12/16/2003 1/13/2004 11/4/2003 12/2/2003 12/30/2003 13,000 175.0010/7/2003 400.00 80.000 Pressure 11,889 168.89 388.89 73.333 10,778 162.78 Overhead reflux 377.78 66.667 9,667 156.67 366.67 60.000 8,556 150.56 355.56 53.333 7,444 144.44 Reboiler steam 444.44 46.667 ~13 Mlb/h 6,333 138.33 333.33 40.000 Tower temperature 5,222 132.22 322.22 33.333 4,111 126.11 311.11 26.667 3,000 120.00 20.000 00:00:00 00:00:00 00:00:00 00:00:00 00:00:00 00:00:00 00:00:00 00:00:00 300.00 P107165 (raw data), psig F107048 (raw data), bpd

FIG. 3

F107050 (raw data), Mlb/h P107482 (raw data), °F

Key operating data.

ronmental regulation enforcement, over-refluxing a distillation tower to ensure that the overhead liquid product specification for impurities is always achieved and avoiding high-temperature constraints for reactor operations, resulting in lower conversion or yield levels. Items 1 and 2 will always result in energy savings.

Item 3 will result in energy savings too, if the increase in reactor performance results in less feed or recycle feed processed.

Here are a few case studies that illustrate some these situations and the energy-related benefits they produced.

Distillation towers. The first case study is applying APC to a fluid catalytic cracker (FCC) gas plant debutanizer tower in an oil refinery. The heating medium for the tower reboiler was 550-psig steam. The APC objective was to maintain the bottom gasoline product Reid vapor pressure and overhead butane product C5 content specifications, while also minimizing energy consumption. The APC was put into service in mid-November and completed by early December. Fig. 3 shows the key operating data during a three-month period. APC lowered the tower pressure and overhead reflux rate (top two trend lines), and due to the lower tower pressure, also reduced the tower temperature to hold the same product compositions. The bottom two trend lines show the resulting reboiler steam reduction and how this reduced the tower temperature by about 9°F. The average reduction in steam rate was about 13,000 lb/hr.

Using steam prices reported previously, and allowing for 10% down-time due to unit or APC issues, the yearly benefit for this steam reduction was about $1,540,000/year at peak prices or $1,020,000/year at average prices. Since this reduction was achieved during the winter months (and a lesser reduction would be achieved during warmer months), half of this benefit is probably more realistic. However, $770,000 per year at peak prices ($513,000/year at average) is still a very attractive benefit.

In another case of distillation towers in a crude unit gas plant, rather than an FCC, reboiler steam reductions were achieved in another debutanizer tower and a gasoline splitter tower. For this project, a post audit was conducted by the control system supplier and the customer. Table 1 shows some of the key operating data outlined in the report. By applying the peak and average steam prices, the resulting yearly energy savings for this portion of the APC project was about $1,090,000 peak ($725,300 average).

These two examples show how lucrative energy savings can be obtained from typical APC applications. Even though APC has been applied to many distillation towers, additional energy savings can be achieved from these applications. The focus of these projects is typically on achieving the product specifications and shifting product yields toward more valuable products. Often there is not enough attention given to reducing the tower pressure to the fullest extent, which would create additional energy savings opportunities.

When energy savings are available on a distillation tower, the tower should be pushed to an absolute limit of overhead condensing capacity, tower flooding, or pumping limits (valve opening).

For many existing distillation APC applications a quick revamp to enhance the pressure minimization function would produce additional and significant energy savings.

FCC reactor/regenerator. Refinery FCC reactor/regenerators

are another good example of where energy savings are commonly achieved. In FCC regenerators, a large amount of coke must be combusted from the catalyst that is continuously circulating between the reactor and the regenerator vessels. A correspondingly Select 165 at www.HydrocarbonProcessing.com/RS

PROCESS/PLANT OPTIMIZATION 4.0 Regenerator flue gas O2 ,%

large amount of air must be injected into the regenerator vessel, which is commonly under 10 to 15 psig of pressure. The blower used to inject this air consumes a substantial amount of energy.

For full-burn regenerators, the operator decides how much air to charge to the regenerator, using an online flue-gas O2 analyzer and regenerator temperatures as the basis of his decision.

The following project results are from a recent FCC APC project. This FCC unit employed a full-burn regenerator design, and a post audit was conducted. Fig. 4 shows the flue-gas O2 data for one year before and one year after APC. The 40% reduction of average O2 levels is an example of how APC can reduce operator conservatism. The APC simply used less air to run the unit.

The resulting reduction in 600-psig air blower turbine steam was 6,280 lb/hr. Again, assuming a 90% process availability and the APC, the resulting yearly dollar benefit was about $742,700 peak ($495,100 average).

Typically, regenerator pressure reduction is also an example of removing operator conservatism. For catalyst circulation to be maintained, a minimum delta pressure between the reactor and regenerator vessels must be assured. Operators will typically stay above this delta pressure minimum, but APC pushes much closer to the actual minimum. It is worth noting that in this particular project, the valves used to control the regenerator pressure were not yet automated and, therefore, the APC could not manipulate the regenerator pressure, as would generally occur. If this manipulated variable had been available, the benefits achieved would have been significantly higher.

Associated FCC gas plant energy benefits were also achieved during this project—about $600,000 peak ($400,000 average).

SPECIALREPORT

3.5

2004 2003

3.0 2.5 2.0 1.5 1.0 0.5 0.0 1 251 June-1

501

751

1,001 1,251 1,501 1,751 2,001 Aug-31 1 hour averages Average and standard deviation comparison

(%) 2003 data period 2004 data period % reduction O2 average 1.16 0.70 40% O2 standard dev.

0.46 0.16 66% FIG. 4

Regenerator flue gas O2 comparison.

The resulting total yearly energy-related benefit was $1,342,700 peak ($895,100 average) for the overall APC project. HP Paul Kesseler has been with Invensys Operations Management as an APC consulting application engineer for 10 years. Prior to this position, he worked at Setpoint/ Aspen Technology implementing multivariable model predictive control technology in a wide variety of process units. Mr. Kesseler has been implementing advanced control applications for the past 23 years. He holds a BS degree in chemical engineering from Texas Tech University. Prior to his APC experience Mr. Kesseler was a process design engineer in the engineering and construction industry.

EXPANSIONS • UPGRADES • GRASSROOTS • RETROFITS • PROCESS OPTIMIZATION

Your InSide Battery Limit Resource C&I ENGINEERING 1930 BISHOP LANE SUITE 800 LOUISVILLE, KY 40218 (502) 451-4977 www.cieng.com

PERSONAL COMMITMENT EXPERIENCE CREATIVITY EXCEPTIONAL VALUE

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HYDROCARBON PROCESSING JUNE 2010

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PROCESS/PLANT OPTIMIZATION

Consider low-voltage AC drives for hazardous areas Such drives used with electric motors can offer significant process improvements and energy savings for equipment used in explosive zones J. RIIKONEN, ABB Oy, Helsinki, Finland

I

gnition of a hazardous atmosphere can happen easily, without warning and with disastrous consequences. Escaping liquid or vapors, in combination with insufficiently protected equipment can cause a powerful explosion, resulting in injuries, damage to the plant, extended downtime and lost production. The ensuing legal claims can take years to sort out. What may have started as a small fault can turn out to have major consequences for the future of the company.

Directives for hazardous areas.

Most accidents can be prevented if safety is made a priority throughout the operation.

The ATEX1 directives have been implemented in the European Union (EU) to promote such an approach. The same principles are increasingly being applied by users all over the world.

Until the advent of the new ATEX Directives from the EU, the installation of electrical equipment in a potentially ATEX 95 The product directive 94/9/EC defines the products/equipment and certification requirements

EHSRs products requirements

explosive atmosphere area (sometimes known as a hazardous area) was a relatively simple procedure, requiring only the selection of equipment certified for the appropriate area. Of course, before ATEX, there were other recommendations, such as IEC standards that were used to ensure safety. But such standards did not provide certification.

The implementation of the new directives changes the perception of the specifier and rotating machine manufacturer, who now become directly concerned. The driven equipment is now specifically covered by the requirements of the directives, as is the compatibility of all the equipment installed on a site.

Two-pronged approach. The legis-

lation has two legs, ATEX 95 and ATEX 137. ATEX 95, the product directive, aims to ensure that equipment for use in hazardous areas is safe and meets applicable standards. ATEX 137, the worker ATEX 137 The worker protection directive 1999/92/EC classifies zones and states the corresponding product category

- Category 1 - Category 2 - Category 3

- Zone 0/20 - Zone 1/21 - Zone 2/22

Zones risk analysis

Profitability, frequency, duration of occurrence of potentially explosive atmosphere FIG. 1

58

Connection between ATEX 95 and ATEX 137.

I JUNE 2010 HYDROCARBON PROCESSING

protection directive, concentrates on the duties of the end user, aiming to ensure that equipment and facilities are managed responsibly when in use (Fig. 1). As of July 2006, organizations in the EU must follow these directives to protect employees from explosion risk in areas with an explosive atmosphere.

While the end user is responsible for the installation and safe operation of equipment, the equipment manufacturer is responsible for the equipment being safe as long as it is operated in accordance with instructions. The end user must define the zone, gas group, temperature class and preferred protection class. Based on this information, the manufacturer proposes the equipment to be used.

Motors and drives in hazardous areas. Electric motors are chosen on the

basis of protection type, which, in turn, is selected on the basis of zones. For instance, flameproof motors are used for Zone 1. It is possible to run hazardous area motors with speed control, as long as the effects on the motor are recognized, mainly to prevent motor surface over-temperature.

1 Directive

94/9/EC provides for the first time harmonized requirements for non-electrical equipment and equipment intended for use in environments, which are potentially explosive due to dust hazards and protective systems.

Safety devices intended for use outside explosive atmospheres, which are required for or contribute to the safe functioning of equipment or protective systems with respect to risks of explosion are also included. This is an increase in scope compared to former national regulations for equipment and systems intended for use in potentially explosive atmospheres.

The ATEX directive consists of two EU directives—ATEX 95 and ATEX 137—describing what equipment and work environment is allowed in an environment with an explosive atmosphere. ATEX derives its name from the French title of the 94/9/EC directive: Appareils destinés à être utilisés en ATmosphères Explosibles.

PROCESS/PLANT OPTIMIZATION The temperature of the motor can be measured directly with thermal sensors in the windings, controlling a certified disconnection. The motor temperature can also be tested by running the motor and lowvoltage AC drive combination across the entire speed range with its load, to ensure that the surface temperature stays within specified limits.

Another way is to control the flow of energy to the motor, thus preventing excess temperature, although this requires some experience of testing. In some industrial drives, the energy flow can be easily controlled through the drive. The motor control platform, direct torque control (DTC), helps control the power transferred and, with that, the motor temperature. It is also possible to purchase an ATEX approved package with motor and low-voltage AC drive.

standards outline the requirements for the different types of equipment. For the end user, the standards relate to the installation, maintenance and repair of equipment in hazardous areas, which has to be carried out according to the instructions from the manufacturer.

Certification by a notified body is required for motors in Category 2. The identifying number of the certifying notified body appears near the CE mark on the nameplate, and the certificate number is

SPECIALREPORT

part of the complementary marking on the motor nameplate.

For Category 3 motors, the directives accept self-certification by the motor manufacturer, based on internal quality control. However, this brings few benefits as the manufacturer is still required to fulfill all the requirements of the standards, and the cost for using the services of a notified body are relatively small in this context. For these reasons, some motor manufacturers use the same certification

Pfl j\\ jk\\c% N\ j\\ jX]\kp%

User selects equipment for zones.

ATEX 137, the worker protection directive, defines the zones for gas and dust, indicating which category of equipment that is to be used in each zone (Fig. 2).

It is the end user’s duty to select the right equipment for the respective zone. The notion of zones does not exist in ATEX 95, the product directive. Manufacturers make equipment for the different categories, and it is the manufacturer’s responsibility to select the right category equipment for each zone.

Zone 0 (called Zone 20 for dust) has a permanent presence of explosive atmosphere and requires equipment of Category 1. No motors can be used in this zone.

Zone 1 (Zone 21 for dust) can have incidental presence of explosive atmosphere during normal duty and requires equipment of Category 1 or 2. Zone 2 (Zone 22 for dust) can have presence of explosive atmosphere by accident, but not during normal duty. This zone can have equipment of Category 1, 2 or 3.

Motor types in Category 2 must be one of the types Ex e (increased safety), Ex p (pressurized), Ex d (flameproof ) for gas, or tD A21 (dust-proof enclosure) for environments with combustible dust. Motor types in Category 3 must be at least Ex nA (nonsparking) for gas, or tD A22 (dust-proof enclosure) for environments with combustible dust types.

?lek\i 9l`c[`e^j `j k_\ nfic[ c\X[\i `e k_\ ?lek\i 9l`c[`e^j `j k_\ nfic[ c\X[\i `e k_\ gif[lZk`fe f] hlXc`kp$\e^`e\\i\[# YcXjk$ [ ] c [ Yc i\j`jkXek df[lcXi Yl`c[`e^j% Gligfj\ Yl`ck ]ifd k_\ ^ifle[ lg# \m\ip ?lek\i Yl`c[`e^ d\\kj Xe[ \oZ\\[j `ek\ej\ jX]\kp Xe[ YcXjk jkXe[Xi[j kf \ejli\ k_\ gifk\Zk`fe c\m\cj pfl `ej`jk fe ]fi g\ijfee\c# Zi`k`ZXc \hl`gd\ek Xe[ gif[lZk`m`kp% :ljkfd ;\j`^e J_fik$ fi Cfe^$K\id C\Xj`e^ G\idXe\ek K\dgfiXip 8ggc`ZXk`fej Dlck`gc\ 9cXjk Fm\igi\jjli\ ;liXk`fe ;\j`^ej Cfn# D\[`ld Xe[ ?`^_ I\jgfej\ ;\j`^ej Gfj`k`m\ Gi\jjli`qXk`fe :cXjj @ ;`m`j`fe ) <c\Zki`ZXc >Xj =`i\ ;\k\Zk`fe JX]\ ?Xm\e :XgXYc\ KiXejgfikXYc\ n`k_ Hl`Zb J\klg 8G@ IG .,)&.,* :fdgc`Xek 9cXjk K\jk\[

J\\ jX]\kp ]ifd fli j`[\ Ç ?lek\i 9l`c[`e^j b\\gj pfl gifk\Zk\[ ]ifd k_\ flkj`[\ `e%

Harmonized standards. To meet

the requirements, harmonized standards exist that manufacturers and end users have to follow. For the manufacturer, the

_lek\iYl`c[`e^j%Zfd

)/(%+,)%0/''

Select 167 at www.HydrocarbonProcessing.com/RS 59

PROCESS/PLANT OPTIMIZATION

Zone 2/22 Abnormal condition Presence of explosive atmosphere only by accident, but not during normal duty (≤ 10 hours per year) Zone 1/21 Occasionally Incidental presence of explosive atmosphere during normal duty (10 to 1,000 hours per year)

Categories 1, 2 and 3

Categories Zone 0/20 1 and 2 Continuously Permanent presence of explosive atmosphere (> 1,000 hours per year) Category 1

CE... II 2 G Ex d/Ex de CE... II 2 G Ex e CE... II 2 G Ex p CE... II 2 D Ex tD A21 CE... II 3 G Ex Ex nA CE... II 3 D Ex tD A22

Combinations for G and D are possible FIG. 2

FIG. 3

Permitted equipment categories and zones.

Equipment manufacturer surveying an electrical installation in a hazardous zone.

procedure as for its Category 2 equipment, with the support of the relevant notified bodies.

If testing is required to ensure a safe motor and low-voltage AC drive combination, this only needs to be carried out once.

Further installations from the same manufacturer can be added without additional testing in similar load conditions.

ATEX and the world. Although

the ATEX directives are only applicable within the EU, they are based on European standards that have led to their principles being demanded by users throughout the world. However, ATEX approval is generally not recognized outside Europe, and industries with hazardous areas need to investigate which approval schemes apply locally.

Select 168 at www.HydrocarbonProcessing.com/RS

The zone system is used all over the world, except in North America, where a system of classes and divisions is used to denote hazardous locations. The classes are based on the type of hazard, while the divisions are based on the risk of explosion that the material presents. Both the EU and the US systems provide effective solutions for managing electrical equipment.

Some US equipment manufacturers have started using zones to be compliant with the ATEX directives, enabling them to sell products and equipment to EU end users. Outside Europe and North America, IECEx is used, which follows the same standards as ATEX. The ATEX directives promote a safe working environment with good housekeeping practices to prevent accidents, along with controls to ensure safely operating equipment. As the ATEX standards gradually take hold over the market, major industrial accidents due to explosions should, hopefully, become a thing of the past. HP

Jari Riikonen is the marketing manager for the chemical, oil and gas division of ABB Oy—low-voltage AC drives—at Helsinki, Finland. He graduated in 1987 from Vaasa Institute of Technology in Finland with a BSc degree in electrical engineering. Mr. Riikoneni has been employed by ABB since 1992.

PROCESS/PLANT OPTIMIZATION

SPECIALREPORT

Trading silicon for carbon: how to reduce energy usage through automation The average plant can conservatively achieve 15% energy savings through this technology D. C. WHITE, Emerson Process Management, Houston, Texas

E

nergy is the largest variable operating cost after raw materials for most of the process industries and its efficient use is key to sustaining profitable operation. Natural gas is the most common incremental fossil fuel and its general increase in price and in price volatility over the past few years are well known with most experts projecting these trends to continue in the future.

Table 1 gives typical specific energy usage (Btu/t product) for common processes1 and the value of a 10% energy reduction in terms of increased financial operating margin at an energy price of $7 per million Btu (mBtu). This value is a significant portion of the total operating margin for most of these processes.

In addition to the direct energy price, it seems likely that the US will eventually adopt some regulations regarding greenhousegas emissions. If the regulations in other countries are a guide, these may take the form of a “cap and trade” on CO2 emissions.

This will place an increased value on energy reductions since these reductions can be used to offset increases in other areas or

can be sold under a cap-and-trade system. If the CO2 reductions are valued at $20/ton, which is at the low range of recent prices in Europe, then reductions in natural gas or equivalent light hydrocarbon fuel usage would have an additional value of approximately $1.3/MBtu. This would add roughly 20% to the value of the energy savings above.

To identify energy savings it is important to understand how energy is normally used. In larger process sites energy distribution can be quite complex (Fig. 1). External purchased energy can be in the form of steam, fuel, electricity or some portion of the raw material that is diverted to direct fuel use.

There can be centralized power or steam production that consumes purchased fuel or raw material. Typically, there will be a variety of users in the process and offsites area including fired heaters; steam and electric drives; and miscellaneous hydraulic, heating and cooling equipment—often supporting distillation/ fractionation operations. Energy recovery in the process can yield

TABLE 1. Typical processes1 specific energy usage Net process energy usage, MBtu/t product Petroleum refining

Value, 10% energy reduction, $/ton

4.4

3.1

29.0

20.3

7.9

5.5

15.0

10.5

Polyethylene

3.3

2.3

Polypropylene

2.0

1.4

10.2

7.1

EO

3.8

2.7

EG

4.5

3.2

Ethylbenzene

2.9

2.1

Styrene

38.8

27.2

Chlorine

25.8

18.1

Ammonia

10.1

7.1

Integrated pulp/paper mill Cement production Chemicals Ethylene (naphtha feed)

EDC/ VCM/ PVC

Process steam drivers

Purchased steam Central steam production Purchased fuel

Central power production Purchased power Central power/steam plant Process/offsites/office FIG. 1

Process steam generated

Process heating/ cooling Process fired equipment

Raw material as fuel

Export steam

Process direct fuel usage

Export fuel Misc. electric usage

Export power

Process electric drivers

Typical industry energy supply/usage sources and uses.

Source: Reference 1

HYDROCARBON PROCESSING JUNE 2010

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SPECIALREPORT

PROCESS/PLANT OPTIMIZATION TABLE 3. Potential energy investments

Plants – highest 10% Average

30% 20%

Plants – lowest 10% Current economic minimum

Potential energy savings

Investment cost (time to implement) Low Medium

High

BTU/ton

Cogeneration Process redesign Replaceing lowefficiency process equipment

Theoretical minimum Medium

General improvements in automation Energy KPI monitoring Operating procedure changes

Low

Increased insulation Steam trap/leak management Enhanced exchanger maintenance Increased condensate recovery

25% Corrected to standard conditions for process configuration, product grades and feedstock quality FIG. 2

Typical process energy saving opportunities.

TABLE 2. Typical North American process industry energy supply/usage distribution

Description

Approximate percentage of total input energy (Equivalent Btu basis, including losses) Chemical Oil Pulp and plants refineries paper mills

Energy inputs 1.

Purchased steam

1

2.

Purchased fuel

63

25

34

3.

Raw material as fuel

(i)

64

42

4.

Purchased electricity

37

10

24

Central utilities 5.

Central steam production

32

34

62

6.

Central Power production

4

5

6

Process energy usage 7.

Process steam heating and steam drives

22

19

39

8.

Process direct fired equipment

27

52

5

9.

Process direct fuel consumption

6

4

10.

Process electric drives

22

15

29

11.

Process miscellaneous electric usage

18

Notes: (i) Included in purchased fuel above

supplemental steam and power production. There will be a variety of miscellaneous electricity users including the office, shop and warehouse facilities. There can also be export of energy to other plants or external consumers.

The relative percentages in each of these areas, based on overall usage, will vary significantly among different types of plants (Table 2).2 The percentages are based on an equivalent Btu basis including losses. For purchased power, this means that “average” primary electricity production inputs of fossil fuels per kWh are considered along with allowances for transmission losses to calculate the Btu requirements rather than just the Btu equivalent of the plant fence-line-delivered power. The process fired equipment percentage includes ambient losses as well as process heat.

Overall, the major energy consumers in the process industries will be the centralized power and steam production, process fired heaters, distillation/fractionation and steam/electric drives.

62

I JUNE 2010 HYDROCARBON PROCESSING

High

Advanced control/ optimization Site energy Management systems Flare reduction programs

For specific plants of the same process type, there is still a wide variation in specific energy usage even after correction to standard conditions for feed quality, product rate/composition and process configuration (Fig. 2). Surveys, including those by the author, repeatedly show wide gaps between the most efficient plants and the least efficient with a 30% spread common between the mean of those with the highest 10% usage and the lowest 10%. This variation is primarily due to the equipment age with older plants often having less-efficient equipment and less heat integration than newer plants. Investments in energy reduction have to be economically justified and the proper level for these investments will depend on the future projected energy costs (including greenhouse gas effects) and the cost of the investments. For each plant there will be a current economic minimum energy use that will vary with projected energy costs. However, there are often additional savings obtainable for even the most efficient plants and 20% additional potential savings can be considered typical for current economics.

Even with these investments most plants would be well away from theoretical minimum energy use which is typically about 25% of current average levels. This would imply that the average plant could conservatively target 15% energy savings to bring its operation to benchmark levels and 35% to reach the current economic minimum.

Those with higher-than-average consumption would have even larger potential benefits. Many plants with comprehensive energy programs have indeed done much better.

A large number of possible investments can be made to reduce energy with differing costs and impacts (Table 3). These range from low-cost/relatively low-impact programs such as reducing steam trap and other steam leaks to installing cogeneration units with very high costs, long implementation times and correspondingly high impact. In this article the focus will be on automation investments. These generally fall in the low- to medium-cost range with savings that are typically midrange. As a result, the expected return on investment for these programs can be quite high. They can also typically be implemented in a relatively short time providing quick payback.

PROCESS/PLANT OPTIMIZATION

SPECIALREPORT

TABLE 4. Potential automation energy saving strategies Reduce process energy demand

Reduce energy supply costs Increase internal utility production efficiency

Reduce external purchase costs

• Advanced control/optimization

• Improve combustion efficiencies—boilers, heaters, kilns

• Energy management system

– Heaters

• Steam management

– Internal vs. external

– Compressors

– Steam header management

– Distillation

(minimize venting, letdown, pressure)

• Improved basic control loop performance

– Boiler allocation

• Maximize process heat recovery/minimize losses – Maximize recovered steam

(electricity generation optimization) – Electricity purchase optimization • Maximize cheaper fuel use

– Blowdown control • Steam vs. electric turbine optimization

• Minimize process recycle (including off-spec. product) • Minimize process pressure drop • Minimize waste/flare gases • Minimize standby equipment • Minimize overdrying • Better low-production rate control

Automation can affect energy use in many ways (Table 4)—more than can be covered in any single article. Overall, they might be considered as a way to trade the relatively small amounts of silicon used in the automation computers and other electronic devices for the large amount of carbon burned to produce energy.

Initially, the strategies can be considered in two categories: those that reduce process energy demand at constant production rate and those that reduce the supply costs. The latter category can be further subdivided into those strategies that increase internal util-

ity production efficiencies and those that reduce external utility purchase costs. Within each of these categories there are multiple strategies that can be pursued. Advanced control and real-time optimization of processes can yield significant savings. Improving basic control loop performance can have a very high payback.

Maximizing heat and steam recovery, minimizing pressure drop, and minimizing waste and recycle are all target areas for automation. Improved control of the steam boilers, power turbines and plant steam system are also likely high-priority areas.

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I 63

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION

Controller PC

Top impurities

LC

Reflux

QC

Top product FC FC

Pump

Measurement Valve

TI

Process Feed flow

FIG. 3

Control temp

Multivariable control

FC

Typical configuration—pump and control valve.

Steam LC

Virtual analyzer Base pump curve

Old design— valve 33% of system pressure drop at max. flow

New pump curve

Bottom QR impurities FIG. 5

Bottom product

Typical two-product distillation column.

Pressure New design— valve < 15% of system pressure drop at max. flow

System friction loss curve—no valve Static head

Max flow

Design flow Flowrate FIG. 4

Reevaluating system pressure drop.

Kenney’s book3 is a good introduction to the subject of process industries’ energy conservation and Shinskey’s book4 on automation energy savings covers many important areas. In this article some selected topics from Table 4 will be covered that, in the author’s opinion, deserve increased emphasis.

Improved basic control loop performance. Efficient

and effective basic control loop execution at the plant is obviously essential to successful plant operation and to other functions, such as advanced control, that are controlling energy. Control loops are composed of a measurement element; an actuator, most commonly a valve; and an executed control algorithm. Improvements in each of these elements can lead to reduced energy usage.

These improvements can often be made very quickly at relatively low investment costs.

Measurement. One of the first areas to evaluate for potential automation changes is improving the measurements of key plant energy-related variables in terms of accuracy, location, frequency and number. Accurate measurement of, and accounting for, energy flows is the first step toward controlling usage.

There are some specific measurement improvements that can have a significant impact. Many plants experience wide variation in fuel gas composition and the corresponding heating value. As discussed in previous papers5,6 mass-based heats of combustion for standard light hydrocarbon fuel-gas components have much less variability than those on a volumetric basis. Even hydrogen, which is one of the major causes of volumetric heating value variability, has a ratio on a mass basis much closer to the other components than its volumetric equivalent. This suggests that fuel-gas control on a mass rather than volume basis will eliminate 64

I JUNE 2010 HYDROCARBON PROCESSING

the effect of much of the composition-based variability and that conclusion is supported by the experience of plants that have adopted such control. With modern flow measuring devices that directly measure mass and also provide a gas density measurement, it is easy to implement such controls.

Hoglen and Valentine 7 discuss a specific example of the improvements obtained with mass measurement on a reformer for hydrogen production. These units may use as feed-site fuel gas, external natural gas or a mixture. Steam-to-carbon ratio control against short-term feed composition variation is important.

Too little steam can reduce catalyst life while too much incurs extra energy costs. In older plants, this control is often done with orifice-plate measurement of the gas flow combined with a gas chromatograph or a mass spectrometer for composition determination. From experience a typical operating margin of 0.2 above the desired steam/carbon ratio target was used to allow for short-term composition fluctuations. Installing direct mass measurement of the feed showed a maximum error of 0.02 in the ratio in tests on the actual plant. A reduction in the operating margin ratio of 0.1 was stated to be worth approximately $500,000 per year in energy savings for an 80-mscd hydrogen plant with fuel valued at $6.50 per mBtu. Pressure-drop reductions. Excessive pressure drop in plant equipment represents unnecessary energy usage. Current operating conditions are often different from those when the plant was constructed or expanded and the plant hydraulic profile is different. In addition, the design may well have been done when energy costs were substantially lower. A systematic critical review of the profile can often reveal many areas where pressure drop and the corresponding energy input can be reduced. The energy savings can be substantial but they are usually composed of many small savings as opposed to a single large opportunity.

There are many pumps in the plant and it is worthwhile evaluating when they are oversized. Fig. 3 illustrates a typical centrifugal pump configuration with a control valve downstream controlling the flow into a process area.

In Fig. 4, a typical hydraulic profile for such a system is shown. For well-designed equipment in the past, the expected system process pressure drop as a function of flowrate would

PROCESS/PLANT OPTIMIZATION

Distillation. It is estimated that there are over 40,000 distil-

lation columns in North America and that they consume about 50% of the energy usage in the refining and bulk chemical industries.8 A typical two-product column is shown in Fig. 5. Improv-

Product value, $/day

High energy cost, $/day

Margin $/day

be estimated, without the valve, and the valve and pump would be chosen to provide a controllable pressure drop at estimated maximum flow conditions—often set at 33% of the maximum flow system pressure drop. The valve type and trim would be set to give good control over the operating range and ideally to linearize, as much as possible, the combined process gain over this range. However, with modern valves, and most particularly valves with modern digital positioners with direct position feedback, it is possible to have precise control at much lower pressure drops.

To obtain the energy savings, it is necessary to reduce the pump impeller size that is normally possible for standard-size pumps.

In the figure the effect of a 15% system pressure drop reduction at maximum flow is illustrated. For a 1,000-gpm pump with an electric drive at 100 psi head operating at 70% overall efficiency, the savings would be $6,000 per year with electricity priced at $0.075 per kwh. This would normally repay the impeller change cost in a few months. In such a review it is, of course, necessary to evaluate startup and shutdown conditions for rangeability and determine if it is also necessary to change the valve trim.

In surveys at a number of plants, we have found many cases where valves are operating at less than 20% open at current maximum plant throughput. These should be a priority target for possible change since the savings can be proportionally much higher than those estimated previously.

SPECIALREPORT

Low energy cost, $/day

Low energy cost margin $/day High energy cost margin $/day

High energy cost optimum

FIG. 6

Low energy cost optimum

Min. reflux high purity specifications

Reflux/ reboiler

Column energy optimization.

ing energy efficiency in this unit operation is obviously an important target area for achieving overall energy savings.

The basic tradeoff is shown in Fig. 6. As reflux and the corresponding reboiler duty are increased, the separation in the column increases and the product value increases. But this increase is not linear. Initially, the reflux has a strong effect on the separation but the effect decreases as more and more is added. Energy costs, however, are approximately linear with reflux/reboiler duty. If the operating margin is then calculated as the difference between the product value and the energy costs, there will be a maximum value and this is the optimum operating point. If energy costs increase

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HYDROCARBON PROCESSING JUNE 2010

I 65

PROCESS/PLANT OPTIMIZATION

1,200,000

600,000

1,000,000

500,000 Extra energy cost, $/yr

Extra energy cost, $/yr

SPECIALREPORT

800,000 600,000 400,000

400,000 300,000 200,000 100,000

200,000

0

0 0

20

40

2

60

3

Excess reflux, % FIG. 9

Excess reflux cost—20,000-bpd stabilizer column.

FIG. 7

7

10

Excess O2 cost—100-mBtu/hr heater.

4

PIC PIC 359 D 357 D

HIC 353 D

5 O2, %

TI 071 AIC 354 D

O2, % CO x 100 ppm

3

Product

Feed FIC 101

H306 TI 072

AIC 356 D AIC 362 D

TI 067

Multivariable control TIC 361

TI 069

FIC 101 TI 073 FIC 101 FIC 101

Combustion control

FIC 361

TI 075

2

1

TI 070

TI 043

O2

CO Pass balance

Fuel balance

Air demand

0 0 min.

FIC 361

FIG. 10

FIG. 8

Process fired heaters

while product values are constant, the margin curve shifts and the optimum moves to a lower reflux/reboiler point. The specific operating targets then should be a function of energy costs rather than a fixed number.

For the case of relatively high-purity specifications in both the top and bottom of the column, a minimum reflux ratio will be required for specific column tray numbers and configuration and feedstock composition. In the author’s experience, many columns operate with reflux ratios far in excess of this minimum required for separation. Fig. 7 shows the extra costs incurred for a standard refinery stabilizer column (20,000 bpd) as a function of excess reflux (steam at $10/mBtu) over the minimum. Note the savings can be hundreds of thousands of dollars per year on a single column. Additional considerations concerning energy use in distillation columns are presented in reference 9.

As illustrated in Fig. 6, it is common to use multivariable control and virtual analyzers for distillation column control, particularly those that are large energy consumers or those with high-purity specifications. A large number of papers have been published on this subject with Blevins et al10 providing a good introduction. Energy-saving results have been very positive with common savings in the 10–30% range.

66

I JUNE 2010 HYDROCARBON PROCESSING

Time

Carbon monoxide versus oxygen in combustion.

Process heaters. Process heaters are major energy consumers

in the refining and chemical industries (Table 2). Many of the opportunities for process heater energy savings are also applicable to steam boilers so proper combustion control is quite important. Heaters and boilers come in many different configurations ranging from very simple package boilers to exceedingly complex chemical cracking furnaces. They can have a single fuel with relatively stable composition or multiple fuels with highly variable compositions. Combustion air and draft control can be natural or forced with multiple control points combined with heat recovery equipment such as air preheaters. Process demand can be relatively stable or highly variable. The prescriptive automation level will depend on the equipment complexity and size with more complicated measurement and control justified on larger and more complex installations.

After ensuring safe and environmentally compliant operation, the next most important control objective is to meet the required heater load demand at the highest possible combustion efficiency, which normally translates into minimum excess-air control. In simple heaters/boilers with stable fuel composition, this is usually translated into air/fuel ratio control based on a load characterization curve. However, as shown in Bussman et al,11 ambient condition changes can have a surprisingly large impact on combustion conditions, enough to move from full to partial combustions conditions under limited ambient changes. When

PROCESS/PLANT OPTIMIZATION fuel composition is varying as well, the problem of proper combustion control becomes even more complicated.

In Fig. 9, the cost of poor O2 control is illustrated for a 100mBtu/hr heater with a 400°F stack temperature rise above ambient and fuel costs at $7 per mBtu. Savings through improved control can be several hundred thousand dollars a year for a single heater.

As shown in Fig. 8, multivariable control is an option and is well-justified for larger heaters with varying loads and varying fuel composition. Flue-gas O2 analyzers are commonly installed on larger heaters with CO analyzers being less common. In the past, CO analyzers were often problematic because of maintenance requirements. However, a new generation of analyzers has appeared with much higher reliability and lower installed costs.

Fig. 10 shows actual CO readings from an industrial heater.

Even with relatively stable O2 readings, there are significant CO excursions indicating incomplete combustion and less than ideal efficiency. To maintain highest-efficiency operation, direct CO measurement and incorporation into the control scheme is recommended.

SPECIALREPORT

site are challenging automation problems. Fig. 13 shows a typical system, though individual sites tend to have unique configurations, particularly older ones that have experienced multiple Reactor section Feed – F

Separation section

Product 1 – P1

Energy to reactor – QS

Product 2 – P2 Energy to separation – QS

Recycle-R FIG. 11

Energy and plant conversion.

Overall plant control. In Fig. 11, a simplified but represen-

Data validation

Margin, $/ton

tational process plant configuration is shown. There is a reaction 150 section with energy input and a separation section, also with $5 per mBtu energy input. In the reaction section, feed is converted to a high100 Optimum valued product, P2, and a low-valued byproduct, P1. They are 50 separated in the separation section along with unreacted feed Optimum that is recycled back to the reactor. Again, typically, as the “per 0 pass” feed conversion is increased, selectivity, which is the ratio -50 of high-valued product to total product, decreases. Process plants with these characteristics include ethane cracking and VCM, -100 among many others. These plants are often operated at conversion -150 targets that are not often changed. However, it is worth evaluating $10 per mBtu whether energy cost changes justify changing plant conditions.

-200 If the energy input in the reactor section is assumed propor-250 tional to the combined feed rate and the amount of combined 0.0 0.2 0.4 0.6 0.8 1.0 feed converted to product, and the energy usage in the separation Conversion section to its combined feed, then it is possible to calculate the FIG. 12 Optimum conversion versus energy. optimum conversion. Overall, as conversion increases, the energy use decreases because of the reduced recycle, but the product values also decrease. The optimum will then depend on the energy Energy management HP BFW Utility RIP system steam cost relative to the conversion selectivity HP vent boilers Coordinated header Process Thermo/ HP coefficient.

Plant data valve control model elec. model import The results are shown in Fig. 12 for Highpressure Process MP HP typical costs, selectivities and conversions. header Power heat recovery export HP/MP Desup.

Details of the equations and coefficients are MP vent HP/MP prv. BFW HP/LP valve provided in the appendix. At a cost of $5 per HP reboilers/ MP exchangers/users import mBtu, the optimum conversion is 38%. As Mediumpressure energy costs increase, the optimum shifts Process LP LP LP/LP header MP reboilers/ heat recovery export prv. Desup. to higher conversions and at $10 per mBtu exchangers/ HP/LP LP vent BFW users valve it is 53%. Not changing the conversion to LP import Lowthe correct value results in approximately a pressure 10% lower operating margin for the higherLP header Process export Cond. to LP reboilers/ energy case. Again note that this change can Condensate condensate BFW exchangers/ treatment Cond. recovery often be made for low or no cost, only a users import Condensate change in the automation setpoints. header Cond. export

Central utility control. Efficient and

effective control of the overall steam and power generation systems in a large process

FIG. 13

Process site utility systems.

HYDROCARBON PROCESSING JUNE 2010

I 67

PROCESS/PLANT OPTIMIZATION

debottlenecking projects and equipment upgrades. There are normally multiple steam headers operating at different pressure levels, each with multiple process users and suppliers. Demand and supply at each level can change with changing plant conditions, and disturbances in one header can propagate to the others.

There is often varying steam import or export to other plants at the site further adding to the disturbances to the overall system. In some plants there will be traditional steam turbine power generation on site as well as newer cogeneration units producing both power and steam.

As discussed further in reference 6, there are four interrelated objectives for steam system automation systems. The first, and highest priority, is to ensure safe operation and react to major and rapid demand/supply disturbances such as a major consumer trip.

The second is to respond to abnormal operations, again generally relating to the supply/demand balance at a particular steam level, and bring the system back to the region of normal dynamic control. The third objective is that of standard dynamic control, i.e., to reject typical disturbances and hold the system close to desired

targets. The fourth objective, and lowest priority, is to operate the system at the economic optimum targets, which generally means meeting desired process utility demands at minimum cost. The overall control system needs to move smoothly between these four objectives and control levels as conditions change.

The figure illustrates the typical advanced controls providing this control. The coordinated header control provides the nonsymmetric action required for abnormal and normal dynamic control while minimizing unneeded steam letdown from one level to the next lower one. Feedforward actions on supply/demand changes are incorporated, often in an overall multivariable control framework. The energy-management system includes plant data capture, validation and reconciliation to establish a reliable base case representing current operation. Equipment models are incorporated into an optimization structure that calculates loading targets that maximize financial margin. These targets are then passed to the lower control levels.

In addition to reduced plant incidents, improved automation in the utilities can generally result in reduced energy usage in the Energy intensity—mBtu/ton—maximum, %

Energy usage—mBtu/hr—maximum, %

120 100 80 60 40 20 0 0

FIG. 14

35

20

40 60 80 Production rate—maximum, %

100

30

90

50 40 30 20 10 0

50 40 30

20

40 60 80 Production rate—maximum, %

20 5

100

B

80 70 60 50

15

40 10

30 20

5

10 0

0 50

FIG. 16

68

55

60 65 70 75 80 85 90 95 100 Base data energy intensity—maximum, %

Specific energy usage histogram, uncorrected (a) and corrected (b).

I JUNE 2010 HYDROCARBON PROCESSING

100

Typical plant-specific energy usage versus production rate.

20 Frequency

60

Cumulative percentage

Frequency

70

10

60

25

80

15

70

90

30

20

80

FIG. 15

100

25

90

0

Typical plant energy use versus production rate.

A

100

10 0

0 50

55 60 65 70 75 80 85 90 95 100 Corrected data energy intensity—maximum, %

Cumulative percentage

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION central utilities area in the range of two to five %—though many plants have observed much larger savings.5 This reference also discusses incorporating site environmental emission limits into the optimization framework—a subject of continued interest.

Analyzing plant energy usage. A key part of any energyreduction program is analyzing current plant operation to determine current average energy usage, the variability in this usage and the reasons for it. In analyzing these data, it is important to bring the usage to standard conditions. In Fig. 14, observed energy use from a typical process is plotted versus production rate. Note that if a linear trend line for the normal operating range is projected back to zero production rate, it does not intersect with zero-energy usage. This is normal. For most equipment there is a sharp rise initially in energy use and then an inflection point where the energy use changes at a much lower slope.

One implication of this pattern is that it is misleading to use an uncorrected energy intensity index number (i.e., Btu/ton product) to benchmark a plant. The observed energy intensity index for the plant is shown in Fig. 15 versus production rate, and the substantial slope of the trend line can be seen.

When evaluating actual data, it is then necessary to correct measured energy intensities to a standard production rate. The first histogram, (Fig. 16a) shows the frequency distribution of the uncorrected energy data. The second, 16b, shows the same data corrected to a standard production rate using the trend line coefficient for the energy usage. The reduction in variability is observable and conclusions based on these data are likely to be more reliable than those on the uncorrected data.

SPECIALREPORT

The same comments apply for feed type and product grades, i.e., corrections are required to bring the data to a standard set of conditions. Generally, this is done by evaluating usage under differing conditions and applying observed differentials. In some locations usage varies significantly with ambient conditions and corrections may be required for seasonal effects, i.e., summer versus winter usage.

When considering an individual equipment item, the evaluation can be more complicated. Consider the energy use for a major steam turbine/compressor combination. With the surge controls there will be significant production range where the energy use is relatively insensitive to rate. No single energy intensity factor, independent of production rate, would adequately model this usage. HP 1

2

3 4 5 6 7 8

LITERATURE CITED Worrell, E. et al, “Energy Use and Energy Intensity of the US Chemical Industry,” Ernest Orlando Lawrence Berkeley National Laboratory Report LBNL-44314, April 2000.

US Department of Energy, Office of Energy Efficiency and Renewable Energy, “Energy Use and Loss Footprints,” http://www1.eere.energy.gov/ industry/program_areas/footprints.html.

Kenney, W.F., Energy Conservation in the Process Industries, Academic Press (1984).

Shinskey, F. G., Energy Conservation Through Control, Academic Press, (1978).

White, D. C., “Advanced Automation Reduces Refinery Energy Costs,” Oil & Gas Journal, October 5, 2005.

White, D. C., “Olefin Plant Energy Savings Through Enhanced Automation,” AIChE Spring National Meeting; April, 2009; Tampa, Florida, Paper 110f.

Hoglen, W. and J. Valentine, “Coriolis flowmeters improve hydrogen production,” Hydrocarbon Processing, August, 2007, pp. 72 – 74.

US Department of Energy, Office of Energy Efficiency and Renewable

Select 171 at www.HydrocarbonProcessing.com/RS

HYDROCARBON PROCESSING JUNE 2010

I 69

SPECIALREPORT

PROCESS/PLANT OPTIMIZATION

Energy, “Distillation Column Modeling Tools,” http://www1.eere.energy. gov/industry/chemicals/pdfs/distillation.pdf.

9 White, D. C, “Energy Use in Distillation Operation: Nonlinear Economic Effects,” AIChE Spring National Meeting, April, 2007, Houston, Texas, Session T8002.

10 Blevins, T.L. et al, Advanced Control Unleashed, ISA (2003), pp. 316 ff.

11 Bussman, W., C. Baukal and R. Sexton, “An Automatic Advantage,” Hydrocarbon Engineering, September, 2008.

12 Sharpe, P., “Optimizing Industrial Utility Plants Within EPA Mandated Limits,” NPRA Computer Conference, Chicago, Illinois. (November 2000) Paper CC-00-142.

APPENDIX Plant model details:

Nomenclature: F —Fresh feed rate, tons per hour R —Recycle rate, tons per hour P1 —Low-valued product rate, tons per hour P2 —High-valued product rate, tons per hour x —Once-through conversion = (P1 + P2)/ (F+R) QR —Net energy input, reactor section, mBtu per hour QS —Net energy input, separation section, mBtu per hour S —Selectivity = P2 /(P1 + P2) CF —Fresh feed cost, $/ton CP1 —Product P1 value, $/ton CP2 —Product P2 value, $/ton CQ —Energy cost, $/mBtu M —Operating margin, $/hr a, b, c, d, e, f, g —Coefficients Equations: R = (1–x)/x

QR = a + (b + c x)(F+R) QS = d + e (F+R) S = 1 – f – gx M = P1 CP1 + P2 CP2 – F CF – CQ (QR + QS ) Coefficients: a — 0.01 b —0.03 c —0.05 d —0.01 e —0.05 f —0.01 g —0.04 CF —600 CP1 —500 CP2 —850 CQ —Case 1—$5/mBtu; Case 2—$10/mBtu

Doug White is a senior principal consultant for the PlantWeb Solutions Group of Emerson Process Management. Previously, he held senior management and technical positions with MDC Technology, Profitpoint Solutions, Aspen Technology and Setpoint.

In these positions Dr. White has been responsible for developing and implementing state-of-the-art advanced process energy automation and optimization systems in plants around the world, and has published more than 50 papers on these subjects. He started his career with Caltex Petroleum Corporation with positions at their Australian refinery and central engineering groups. Dr. White holds a BS degree from the University of Florida, an MS degree from the California Institute of Technology, and MA and PhD degrees from Princeton University, all in chemical engineering.

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I JUNE 2010 HYDROCARBON PROCESSING

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SAFETY/LOSS PREVENTION

Performing correct initial blowdown calculations These guidelines can help determine peak flowrate A. VYAS, Sharjah, United Arab Emirates

T

he primary function of blowdown is to provide equipment with the means to reduce pressure when containment loss is unacceptable. The reduction in pressure compensates for the increase in temperature that takes place during fire conditions such that the coincident pressure and metal temperature do not exceed the allowable stress on equipment and piping. A secondary function of the blowdown system is to reduce local containment loss that arises from leaks that may otherwise lead to escalation and the risk of catastrophic structural failure.

A consequence of blowdown is the attaining of low temperature for process fluids and equipment. Consideration must be given to the auto-refrigeration effects caused when a vessel is depressurized.

These effects increase when the vessel is depressurized when there is no fire present.

Consideration must also be given to the temperature that will be attained by the blocked-in system if left to stand after isolation prior to initiating blowdown. Such conditions set the lower design conditions of the system. Blowdown is initiated remotely by the operator via the emergency shutdown system, which will allow blowdown to commence after isolation via the emergency shutdown valves (ESDVs).

Blowdown zone segregation and staggered blowdown. The total

number of blowdown zones is based on flowrates and area. Blowdown staggering is determined if a fire occurs in any one of the fire zones that needs to be depressurized, immediately followed by the remaining fire zones to be depressurized in the controlled sequence.

All blowdown sections can be depressurized in a controlled sequential staggering during an emergency situation

without exceeding flare capacity. The total maximum depressuring time has to be estimated, usually about four hours, to empty out complete system inventory up to atmospheric pressure, considering that all blowdown sections are at normal depressurization rates. The staggering sequence depressuring details for all blowdown zones during fire has to be calculated in detail.

Repressurization. Adiabatic depressurization will lead to low temperatures in the equipment. Once the depressurization is completed, immediate repressuring of a system from minimum design temperature could affect the mechanical integrity of the systemâ&#x20AC;&#x2122;s construction material. The possibility of immediate repressurization should be avoided by a timer device that will prevent immediate repressurization, hence, coincidental high pressures and low temperatures should be avoided. As a rule, repressuring should be carried out in specific steps. Also, you must consider the low temperatures generated by the Joule-Thompson effect across the repressuring valve as the respective blowdown system pressure will be close to atmospheric pressure.

Coincidental conditions. After com-

pleting the adiabatic depressurization, immediate repressuring of the system from minimum design temperature may affect the mechanical integrity of the system construction; therefore, coincidental high pressure and low temperature should be avoided. The repressuring step time after depressurization shall be maintained such that the membrane stress limits do not exceed the 50N/mm2 design conditions as per standard.

Emergency depressurization. Plant emergency depressurization can be exe-

cuted in two ways: zonal depressurization or total area depressurization.

Zonal depressurization. This is

executed based on manually selecting particular blowdown zones in the event of any emergency or fire that depressurizes the selected blowdown zone. During the manually selected zonal depressurization, all other zones will be inhibited from depressurizing until a determined time delay has elapsed to ensure that the instantaneous capacity of the flare system will not be exceeded. Once the preset time delay has elapsed, another zone would then be allowed to be depressurized if selected.

Total area depressurization. This

is executed to depressurize an entire area in sequential order or staggering by opening all blowdown valves in the quickest amount of time within constraints of the flare capacity. Time delays, as defined, are based on the premise that the second blowdown zone is the largest remaining depressurization load. This ensures that the instantaneous load on the flaring system following the time delay will not exceed the flare capacity.

To reduce the overall area depressurization time, the blowdown zones are to be interlocked such that depressurization of the zones is staggered. The time between initializing the depressurization of a zone, during total area depressurization, has been arranged so that the flare capacity is never exceeded. Upon activation of total area depressurization, each section of the area will be depressurized based on staggering. All blowdown sections can be depressurized in a controlled sequential staggering during an emergency situation without exceeding the flare capacity.

HYDROCARBON PROCESSING JUNE 2010

I 71

SAFETY/LOSS PREVENTION a high point without a low point between the concerned equipment or line and the blowdown device. On equipment (drum, column, etc.) having a gas device (e.g., mist eliminator), they shall be installed upstream of the gas device. It is a common practice to fit the BDV with limit switches indicating in the control room that the valve has actually moved to its required position.

A standard describes two emergency depressuring criteria that have historically been used by the petrochemical industry.

The criteria of depressuring to only 50% of design pressure within 15 mins may be suitable for pool fire exposure of vessels with a wall thickness greater than 25 mm.

But it may not be suitable for thinner vessels, for torch fire exposures, or fire exposure of piping.

The use of vessel/piping passive fire protection (e.g., fire-resistant insulation or fireproofing) may reduce the temperature-rise rate of the fire-exposed equipment. The analysis required to accurately determine the survivability of equipment exposed to a torch fire can be complex.

This dynamic analysis involves defining the heat flux, calculating the heat-up time

Disposal system components design. Depending on the process plant

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under consideration, a disposal system may consist of a combination of the following items: piping, knock-out drum, quench drum, seal drum, flare stack, ignition system, flare tip, and burning pit.

Blowdown device. The blowdown

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device concerns the automatic or remote blowdown and there are two valve types: • A full bore ball valve that fails open and is specified as “tight shut-off ” with a minimum size of 2 in. and a restriction orifice installed downstream the blowdown valve (BDV) and sized for the required flowrate.

• An on/off control valve type sized for the required flowrate with a mechanical opening limiter stop.

A reducer is installed downstream of the restriction orifice or control valve type BDV followed by a manual block valve full bore, ball valve, CSO having the flare sub-header line size. The installation of elbows between the restriction orifice and the reducer is forbidden. The blowdown device shall always be installed on

Antisurge valve Flare

Flare PSV

Process stream

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PSV Length

P: 281.987.0726 F: 281.987.0905

Hypothetical segment FIG. 1

Process stream

Modeling the hypothetical segment used for sizing procedures.

SAFETY/LOSS PREVENTION of the vessel/piping walls, changing wall stresses and reduced strengths. The impact of high-rate depressuring on vessel internals and/or a catalyst should be appropriately considered when establishing the depressuring rate.

Reliability requirements. The depressuring system design should consider features to reduce spurious trips. Secured air supply systems should be specified for each depressuring system to prevent the common-mode opening of emergency depressuring valves.

If sequenced depressuring is adopted, the system design shall ensure that: • System failure cannot result in an uncontrolled simultaneous depressurization of the whole facility • System failure cannot result in a situation where automatic depressurization of the installation is prevented.

THREE OPTIONS IN EMERGENCY DEPRESSURING VALVE DESIGN IMPACTING RELIABILITY 1. Air fail-open valve with normally energized (NE) solenoid.

• If there is an air failure, the valve opens. Spurious trips can be avoided using secured air buffer vessels.

• If there is a power failure, the solenoid will vent and the valve will open. Spurious trips can be avoided by designating the power source to be uninterruptible power supply (UPS). However, signal cable failure will result in a spurious trip.

• This design option is the most common for high rate emergency depressuring.

• See 1 or 2 for issues regarding NE or NDE solenoids.

• This design option is typically only for low-rate operational depressuring. The use of Option 3 for high-rate emergency depressuring requires technician approval.

The reliability of the system shall include upstream equipment analysis during and after emergency depressuring. The upstream system protected by emergency depressuring valves shall be designed to cope with peak depressuring (high velocities) and with the final end-process conditions (very low process temperatures that might be generated when depressured).

in most cases to low temperatures (often well below 0°C) within the equipment and its related piping. This has an impact on selecting the equipment material. After emergency depressuring, operational steps to reestablish normal operation may lead to process conditions that might not have been considered in the design. For example, quickly restarting the depressured equipment without letting the equipment warm up first, or using other process/utility streams to accelerate the start up, may lead to higher stresses than the material can safely withstand at cold temperatures.

Required details to start blowdown calculations - Description of the fire /adiabatic scenarios - Blowdown sectionalization (volume inventory) - Follow standards for design basis and isolation philosophy - Manual or automatic blowdown - Acceptance criteria for rupture Step 1 Calculate the size of all orifices and all peak rates segments.

Reduce the orifice size.

Is the flare system capacity utilized when adding all the simultaneous blowdown rates together?

No

Evaluate the blowdown rate, preferably for the most hazardous blowdown section.

Step 2 Check the thickness of the vessel as well as pipe.

Add insulation if required.

Yes

No

Is the blowdown rate less than maximum? l- dP/ dt l

Step 3 Yes Calculate the inner-wall temperature profile for all pipes and equipment. Use the local fire with the highest heat flux (kW/m2).

Step 4 Is acceptance criteria for rupture met?

2. Air fail-open valve with normally de-energized (NDE) solenoid.

Yes

• If there is an air failure, the valve will open. Spurious trips can be avoided using secured air-buffer vessels.

• With the NDE design, spurious trips due to power loss are avoided. Unless there are alternate means to vent the emergency depressuring actuator (e.g., manually venting the air supply at a field panel), NDE design use requires measures for reliable functioning of the depressuring valve during a fire.

No

Step 5 a) Decide which pipe/equipment to fire insulate, or b) Increase orifice diameter if there is available capacity in the flare system, or c) Reduce system volume by relocation of sectionalization valves, or d) Increase the flare system capacity, or e) Change material quality, or f) Increase wall thickness

Step 6 Calculate the minimum design temperature (low-temperature design temperature) of the blowdown section and the flare system tailpipe.

Is the minimum design temperature acceptable?

No

Check flexibility again with basis of design calculations. Check ambient conditions. Check with piping whether low temperature limit is acceptable for low cost material or not.

Yes

3. Air fail-close valve with NE or NDE solenoids.

• If there is an air failure, the valve will remain closed. Since air system failures during fire exposure may be possible, fire protecting the air system is required.

Brittle fracture. Depressuring leads

Finish the RO design with calculated peak rate and material selection of upstream and downstream of BDV.

FIG. 2

Simplified chart to perform initial blowdown calculations to find orifice size and minimum design temperature.

HYDROCARBON PROCESSING JUNE 2010

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SAFETY/LOSS PREVENTION The construction of equipment materials shall be selected accordingly.

The specified lower design temperature for the piping and equipment shall take into account the low temperatures that can be experienced by equipment and piping TABLE 1. General guideline for process disposal Material

Vent Flare Process Sewer

Process vapors Flammable, non-toxic and toxic

X

X

X

X

both upstream and downstream of a depressuring device when depressuring.

Material selection. Materials shall

be selected taking into account how the systems are sectionalized and depressured.

The temperature of process fluids should be calculated as a function of pressure for each depressuring application. These temperatures shall be compared to the vessel and piping minimum safe operating temperature to determine if any restrictions shall be placed on the depressuring or material selection altered.

Process vapors Nonflammable and toxic

Software modeling—real process segment used as a hypothetical situation. Fig. 1 shows the minimum

Process vapors Nonflammable and nontoxic

X

Steam

X

Sewer vapors

X

X

Liquids Process blowdown

X

Thermal relief

X

X X

Process drain

X

Surface runoff

X

requirements for performing proper depressurization and fire-adiabatic calculations together with a procedure for sizing depressurization and relief systems for pressurized systems exposed to fire with the use of computational dynamic depressurization.

Before proceeding with the fire model, some terminology will be defined.

Modeling process segments. A

model (Fig. 1) is represented for real geom-

HIGH ACCURACY FLOW METERS

Software and analysis. After a careful

check/analysis of input data, the depressurization study can be started using a computational dynamic depressuring study for fire and adiabatic cases to determine the peak flowrate and minimum temperature for the individual systems, respectively. Fig. 2 illustrates how to perform an initial blowdown calculation to find the correct orifice size and minimum design temperature. HP

FOR HIGH TEMPERATURES AND HIGH PRESSURES – non-intrusive ultrasonic clamp-on technology – for temperatures up to 750 °F – independent of process pressure – multi-beam for high accuracy – wide turn down – installation without process shut down – no maintenance – no pressure loss – standard volume calculation

www.flexim.com usinfo@flexim.com FLEXIM AMERICAS Corporation Phone: (631) 492-2300 Toll Free: 1-888-852-7473

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Select 174 at www.HydrocarbonProcessing.com/RS 74

etry of one complete isolated segment that was protected with a relief and blowdown valve. Segments that represent the total system volume and heat-transfer areas have different wetted areas and wall thickness. The hypothetical segment is used for calculation of the system pressure during depressurization or relief. Sectionalisation valves close in emergency situations and blowdown valves open in fire situations.

Notice that the hypothetical segment is modeled with a real system volume, system outer-area without actual system inside-area in contact with the gas, system inside-area in contact with the liquid. The hypothetical segment is modeled as a cylinder. This shape is specified with a diameter equal to the most dominating (volume) diameter of any piping item or equipment in the original segment. The associated wall thickness for this diameter should be used. The cylinder length is set such that its volume equals the original segment’s volume. The liquid level is adjusted to obtain the actual liquid volume. The hypothetical cylinder now represents the correct system volume. However, the outer area of this cylinder must be corrected to the real system outer area by trialand-error for the addition or subtraction of a gas and liquid area to the segment.

BIBLIOGRAPHY American Petroleum Institute, “Guide for Pressure-Relieving and Depressurizing Systems,” API Recommended Practice (RP) 521, Fifth addition, January 2007.

Vyas, A., “Determining system depressurizing requirement,” Hydrocarbon Processing, pp. 89–90, November 2008.

Salater, P., V. Overaa and E. Kjensjord, “Best Practice on Depressurization and Fire-Relief Design,” Norsk Hydro ASA. Nolan, D. P., Handbook of Fire and Explosion Protection Engineering Principles for Oil and Gas.

Archana Ajay Vyas is a process engineer with seven years in the oil and gas industry. Her areas of expertise are flare network system design and depressurization system calculations. She earned a BE degree in chemical engineering from GECG, University of Gujarat, India.

PLANT DESIGN AND ENGINEERING

How to avoid interface engineering problems Here are several ways to accurately exchange information S. MOULIK, MWKellogg, Ras Laffan, Qatar; K. SAITO, Chiyoda Corporation, Yokohama, Japan

I

nterface engineering is part of the detailed engineering of plant design when two or more engineering, procurement and construction (EPC) contractors are involved. It deals with the exchange of design, schedule and construction details between all contracting parties.

In the past, a large plant used to be built by one EPC contractor. Today, installations are being put up in a much bigger scale to make it more cost-effective. The time frame from when the investment decision has been made to commissioning is also getting shorter. To avoid any schedule or budget overrun risk, the owner may divide the whole complex into several smaller contracts implemented by multiple EPC contractors. The timely and accurate exchange of design, schedule and construction information between design offices becomes important. This exchange of information is to minimize rework at a later date and avoid last-minute surprises during commissioning.

It becomes the project management team’s task to act in advance to identify the interfaces before the packages go for bidding. In case the interfaces are not identified or not identified properly, then the detailed engineering contractors need to depend heavily on the accuracy of front-end engineering design (FEED). Design development/changes are normal between FEED and detailed engineering. These changes may not be captured in the integrated design until unless detailed engineering design information are exchanged between two or more parties at contract boundaries, consolidated and subsequently reconfirmed. Shortcomings in interface identification typically lead to inadequate flare design, existence of the same line in two different sizes in two contract boundaries, missing lines, pipe elevations mismatching at expected match lines, clashing of one contractor’s cable trenches with other contractors under ground pipe, etc. Mostly, it is the utility systems and product/ raw material systems that get adversely affected by inadequate or no exchange of design information.

Identification of interfaces is a critical task. It needs involvement of all engineering disciplines to identify their respective discipline interfaces between different unit/project boundaries.

Construction engineers also need to be involved in identifying the construction interfaces, e.g., welding and hydro test sequence, scheduling between two contract boundaries, etc.

During an interface identification period, there is always a probability to miss an interface. The detailed engineering contractor should recognize these potential shortcomings early

during the detailed engineering phase. Early identification of interfaces also helps reduce the increased scope that may occur if interfaces are not adequately recognized.

Identification of interfaces is normally being done by the FEED contractor or by an independent project management team. As this identification work can be executed under one roof, it may be termed as manageable. However, to make agreements between two unwilling EPC contractors, working from two different geographical locations, requires a completely different approach. This becomes the interface engineer’s skill and expertise to resolve the issue strictly under project guidelines without extra costs. The responsible person from the owner side requires strong leadership quality and drive, as well as good coordination skills, to manage changes arising from the interface.

In interface engineering, the first step is to identify interfaces.

The second step is to get the parties to agree on content and scheduling the information exchange. The third step is to ensure that exchanging proper design information occurs in the time frame agreed. Finally, the most important task is to ensure that the interfacing parties agree upon each other’s information.

Exchanging information may take place using normal e-mail, hard-copy exchange or through a secured Web page—whatever suits the project most. Traveling from one design office to another, to resolve conflicts, is usually limited. While exchang-

FIG. 1

Piping interface showing two different contractors’ pipe matching in a 3D model.

HYDROCARBON PROCESSING JUNE 2010

I 75

PLANT DESIGN AND ENGINEERING ing information remotely is effective, it has its limitations. On a big project, a dynamic situation occurs when all design offices start work simultaneously. In this situation, working in relative isolation without significant interaction with other parties may become very expensive and may impact the project schedule. It becomes necessary to plan regular discussions between design engineers of two respective design offices. Audio/video conferences with monthly/bi-monthly face-to-face meetings minimizes mistrust and encourages team building. However, there is no alternative to face-to-face meetings and payback in terms of com-

FIG. 2

Clash in interface area identified after loading one contractor’s 3D model data in another contractor’s model.

Facilities redesigned to make it clash free.

pleting the design more than outweighs any of the expenditure incurred for traveling.

Today, plants are modeled in 3D. It is desirable that all EPC companies working on the same project use the same platform so that the design files of the interface area can be exchanged and uploaded in each other’s design without any additional effort (Fig. 1). However, this is not always achievable since the contracting design offices have developed their own respective packages over a period of time. Also, changing to a different software package is not cost-effective for a single project.

When the time comes for exchanging 3D interface design files, several people franchise their veto power. They claim that the company may lose vital proprietary design information through interface design data exchange. This argument is not correct. Design information of an interface area is known to both interfacing parties; hence, exchanging the 3D model of the interface area does not reveal anything new and certainly won’t harm the project.

Exchanging 3D model information does not only confirm the accuracy of the exchanged information but it also helps one contractor to check clash in another contractor’s area (Fig. 2).

Here are a couple of common questions and answers concerning engineering interfacing: Q: What if the contractor does not provide information on time?

A: This situation will not arise if there is a synergy in all the contractors’ project schedules. It’s the owner’s responsibility to ensure that the project awarded to different EPC contractors have similar project schedules and that they are realistic.

Milestone payments may be linked with the important interfacing deliverables to ensure that contractors perform in a timely manner. Keeping a continuous vigil on the critical interfacing issues is effective and avoids unnecessary clash between two interfacing teams.

Q: What if one interfacing party does not agree with the other’s deliverables?

A: This is common. The interface engineer must ensure that the documents are being reviewed strictly against project specifications. Ensuring that both interfacing parties follow the same project guidelines or agreement, then a disagreement should not happen.

Remember—proper communication between two contractors is the most important aspect of interface engineering and that is the real challenge. HP

Subhendu Moulik is an associate project manager for MWKellogg in the UK. He received a BE degree in mechanical engineering from the National Institute of Technology in Surathkal, India. Mr. Moulik is a member of the Institution of Engineers in India and also a member of the Project Management Institute in the US. He has worked for the past 17 years in different types of process industries that include chemical, petrochemical, refinery, and oil and gas. Mr. Moulik started as a maintenance engineer and progressed his career through construction engineering and project management.

Koichi Saito is an interface coordinator for Chiyoda Corporation, Japan. He received a BE degree in mechanical engineering from the Institute of Science and Technology at Nihon University, Japan. Mr. Saito has worked for 21 years in the natural gas industry, and the fertilizer industry and has been a project engineer for an industrial plant.

Select 175 at www.HydrocarbonProcessing.com/RS

HEAT TRANSFER/RELIABILITY

Troubleshooting waste-heat boiler poor heat transfer Analytical results showed that potash leached from the catalyst, refractory powder, corrosion products and, to a lesser extent, catalyst fines resulted in tube fouling G. YEH, I. AL-BABTAIN and S. AL-ZAHRANI, Saudi Aramco; and N. AL-GHANEMI, PetroRabigh

O

ne of Saudi Aramco’s plants encountered a steady process temperature rise at the waste-heat boiler (E4) outlet of the hydrogen-producing unit (HPU) after an upset that occurred in February 2008. The upset initially resulted in unsteady firing in the steam-reformer furnace and lower reformer outlet temperature. The feed rate was cut with the intent of bringing up the reformer outlet temperature. During the upset, the reformer outlet temperature temporarily increased to 850°C and the reformer inlet temperature dropped to 320°C. These temperature upsets did not impose any ill effect on the reformer performance. The steamdrum level, steam-reformer inlet and outlet pressures, the liquefied petroleum gas (LPG) feed rate and the steam/carbon (S/C) ratio were in an appropriate range during the upset.

The HPU experienced a continuous process temperature increase at the E4 outlet since the upset. The normal process outlet temperature is 350°C, but it increased to 363°C in May (Fig. 1).

The E4 outlet temperature subsequently decreased to 360°C after the LPG feed rate and the S/C were reduced. The E4 process outlet temperature is the inlet temperature of the high-temperature shift (HTS) reactor, R2. The corresponding HTS outlet temperature

increased from 397°C to 409°C in May (Fig. 2). The maximum design temperature for the HTS converter is 450°C. The continuous process temperature increase at the E4 outlet not only increased the HTS converter inlet and outlet temperatures, but also increased the desulfurizer, R1, inlet temperature (Fig. 3). The desulfurizer inlet temperature reached 398°C in May. The desulfurizer has a 400°C maximum allowable temperature above which coking will form and deactivate the CoMo catalyst in the desulfurizer. Once the CoMo catalyst is deactivated, organic sulfur can pass through the desulfurizer and poison the steam-reforming catalysts. A simplified process flow diagram of the HPU is shown in Fig. 4.

The E4 outlet process temperature increased continuously after the upset, while the LPG flowrate, reformer outlet temperature and the S/C ratio remained fairly constant. This phenomenon indicated that the E4 heat duty was decreasing. The calculated E4 heat-transfer coefficient decreased while the E4 process outlet temperature increased (Fig. 5). Root causes of E4 heat-duty loss can be: • Internal bypass valve opening • Tube-side fouling • Shell-side fouling.

365

420

360

410

FIG. 1

E4 process outlet temperature.

FIG. 2

6/1/08

5/12/08

4/22/08

12/24/07

6/1/08

5/12/08

4/22/08

4/2/08

3/13/08

2/22/08

360 2/2/08

330 1/13/08

370

12/24/07

335

4/2/08

380

3/13/08

340

390

2/22/08

345

400

2/2/08

350

1/13/08

Temperature, °C

Temperature, °C

355

HTS outlet temperature.

HYDROCARBON PROCESSING JUNE 2010

I 77

HEAT TRANSFER/RELIABILITY Internal bypass valve opening. E4 is equipped with an

The plant personnel stroked the bypass valve and found the valve worked properly. The field check by inspection shows that the distance between the flange and connection point (actuator indicator) is 410 mm; nevertheless, the drawing shows the distance should be 405 mm for the fully closed position. The bypass valve might be pushed in 5 mm more than the design indicates.

The bypass should be checked during turnaround/inspection (T/I) for this discrepancy.

During a shut down to repair the leaking thermowell in October 2008, the plant found that the bypass valve was in good working condition.

internal bypass valve to regulate the process temperature. The internal bypass valve opening can increase E4 outlet process temperature. The effect of the bypass valve opening should be a stepwise outlet temperature increase, and after a stepwise increase, the process outlet temperature should stabilize. However, the HPU experienced a continuous outlet temperature increase. If the bypass valve opening was the root cause, the opening must be enlarged gradually to cause a continuous outlet temperature increase.

In a test run, the reformer outlet temperature was increased from 805°C to 815°C and the E4 outlet temperature increased 2°C to 3°C. According to the calculation, if the internal valve opening is the root cause, the E4 outlet temperature should only increase by 0.3°C. This test result does not support that the bypass valve is either open or closed.

Tube-side fouling. The E4 outlet temperature’s steady increase

405 400 395

Temperature, °C

390 385 380 375 370 365 360 355

FIG. 3

6/1/08

5/12/08

4/22/08

4/2/08

3/13/08

2/22/08

2/2/08

1/13/08

12/24/07

350

Desulfurizer inlet temperature.

R-1

R-2 Hydrogen EX PSA unit

Stack

V-7

since the upset implies that the problem might be on the process side. Tube-side fouling can be formed by catalyst fines, refractory powder, corrosion products, and potash and silica deposits. The catalyst vendor indicated that it is unlikely that fouling is caused by catalyst fines, and the catalyst, when broken, will break into a few big pieces instead of fines. That the steam-reformer pressure drop (Fig. 6) remained constant, with relatively constant LPG and steam rates, confirms the catalyst vendor’s opinion.

Infrared readings indicated that the E4 inlet cone skin temperature is around 285°C. This finding indicates that the refractory on the cone shape of the E4 is not damaged. The HPU uses the potash-promoted steam-reforming catalysts required to prevent coke formation. In normal operating conditions, potash is released slowly to prevent coke formation; however, potash can be released faster than necessary during a high-temperature upset and can migrate to the E4 and foul the tube side. A large potash release can cause coke formation, hot tubes and high-pressure drops in the reformer tubes. Nevertheless, the steam-reformer pressure drop was normal and there were no hot spots on the reformer tubes indicating that a significant amount of potash released from the current batch of catalyst is unlikely.

It was found that refractory on the E4 cone head contained 35%–39% of silica and there is a stainless-steel (SS) lining to protect the refractory from silica leaching. The SS lining can be damaged due to an upset, which Export MP steam can cause silica leaching. Silica fouling on the tube side might be one of the root causes for the high E4 outlet temperature. We recommended that the plant adopt the new dual-refractory system design (insulation layer of silica-alumina and hot-face layer of pure alumina) instead of using the SS lining to prevent silica leaching in the next T/I. The HPU was shut down to repair a leak in the thermowell in October. Subsequently, the tube side was found fouled.

LPG feed

Shell-side fouling. The water quality from F-1

E-4

Gas to coolers Condensate

FIG. 4

Process flow diagram.

Deaerator

Demineralized water

F-1: steam reformer R1: desulfurizer R2: HTSC E-4: wasteheat boiler

the steam drums, i.e., phosphate, pH, iron and total dissolved solids, was generally good from January up to May. The steam drum level was maintained between 57% and 62% during the upset. The plant does not have a history of water fouling in the E4. All the above indicate that water-side fouling is unlikely.

Reformer and HTS reactor integrity.

It seems the reformer was working fine and was 78

I JUNE 2010 HYDROCARBON PROCESSING

320

365

310

360

300

355

290

350

280

345

270

340 Heat transfer coeff Process outlet T

260

FIG. 5

2.5

335

2.0 1.5 1.0 0.5

E4 heat transfer coefficient.

FIG. 6

not affected by the upset. Pressure drop, tube-wall temperature (TWT) and methane slip (Fig. 7) are in an appropriate range. A little higher methane slip during May 2008 can be attributed to lower S/C ratios and reformer outlet temperatures.

HTS converter pressure drop was stable at around 0.5 kg/cm2 and the carbon monoxide slip (CO%) at the HTS outlet was quite steady (Fig. 8). CO slip can be affected by the HTS inlet and outlet temperatures. The HTS should not be operated at unnecessarily high temperatures that will shorten the catalyst life.

6/1/08

5/12/08

4/22/08

4/2/08

3/13/08

2/22/08

2/2/08

1/13/08

0.0 12/24/07

5/12/08

4/22/08

4/2/08

3/13/08

2/22/08

2/2/08

330 1/13/08

12/24/07

250

3.0

Pressure drop, kg/cm2

370

E4 process outlet T, °C

330

6/1/08

U, Kcal/m2 °C hr

HEAT TRANSFER/RELIABILITY

Steam reformer pressure drop.

High E4 process outlet temperature effect. A high E4 process outlet temperature causes a high desulfurizer inlet temperature. The desulfurizer inlet temperature reached 398°C in May 2008. The unnecessarily high temperature at R1 deactivates the catalyst faster and should be prevented. The desulfurizer has a maximum allowable temperature of 400°C, above which, coking will form and deactivate the CoMo catalyst in the desulfurizer.

Once the CoMo catalyst is deactivated, organic sulfur can pass through the desulfurizer and poison the steam-reforming cata-

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HYDROCARBON PROCESSING JUNE 2010

I 79

HEAT TRANSFER/RELIABILITY TABLE 1. Foulant sample analytical results from the catalyst vendor and the plant

6.0 5.5

Methane slip, %

5.0 4.5 4.0 3.5 3.0 2.5

FIG. 7

6/1/08

5/12/08

4/22/08

4/2/08

3/13/08

2/22/08

2/2/08

1/13/08

12/24/07

2.0

Methane slip.

5

CO slip, %

4

Catalyst vendor data, % 32.09 25.60 0.12 0.10 0.18 19.03 3.09 0.25 0.91 0.37 14.90 1.17 0.26 0.20 1.17 0.05 0.39 ND ND 0.32 99.62

Plant data, % 20.70 20.53 ND ND 1.05 33.68 1.93 0.18 1.05 0.35 14.74 2.11 0.18 ND 0.70 ND 0.53 1.93 0.18 0.35 100.18

ND: Not detectable or level below detection limit.

3

minimum S/C ratio should be set at 3.8 to prevent coke formation in the steam reformer. Meanwhile, maintain the recycled H2 to the LPG molar ratio above 12% to minimize coke formation in the desulfurizer.

2

FIG. 8

CO slip at HTS outlet.

FIG. 9

Foulant sample.

6/1/08

5/12/08

4/22/08

4/2/08

3/13/08

2/22/08

2/2/08

1/13/08

12/24/07

1

lysts. To prevent deactivating the CoMo catalyst, the LPG feed rate, S/C ratio or reformer outlet temperature should be reduced to maintain the desulfurizer inlet temperature below 400°C. The 80

Al2O3 SiO2 P2O5 SO3 Cl K2O CaO TiO2 Cr2O3 MnO2 Fe2O3 NiO ZnO Ga2O3 ZrO2 Nb2O5 MoO3 Na2O MgO PbO Total

I JUNE 2010 HYDROCARBON PROCESSING

Tube-side foulant sample. In October 2008, the HPU was shut down to repair a leak in the thermowell that measures the reforming process outlet temperature. Subsequently, it was found that the E4 bypass valve was in good condition, but the tube side was fouled. The plant did not open the shell side since it found foulants in the tubes and did not expect fouling on the shell side.

Foulants deposited at the bottom of the tubes were powdery materials. The foulants on the surface are hard and difficult to remove. No hydrocarbons were found in the foulants. The foulants were black and gray (Fig. 9). Boroscopy showed that refractory on the tube sheet at the E4 inlet is slightly damaged and refractory on the cone shape of the E4 and the SS lining were in good condition.

The foulant sample analytical results reported by the catalyst vendor and the plant are summarized in Table 1.

Analytical data from the plant and catalyst vendor were consistent. The fouling materialsâ&#x20AC;&#x2122; source was identified as potash leaching from the steam-reforming catalyst, refractory powder, corrosion products and catalyst fines.

This potash in the foulant sample did not leach from the current batch of the steam-reforming catalysts since the current batch is performing well. This high potash level found in the foulant sample is due to a cumulative effect over many years during which many new catalyst charges have been installed and many steamings have been conducted to reduce the pressure drop.

To protect the equipment, and have a long operating cycle, it is critical to have stable operation in the hydrogen plant. Potash in the steam-reforming catalysts will release gradually to avoid coke

HEAT TRANSFER/RELIABILITY formation in normal operation and it will release at faster rates during steaming and high-temperature conditions. Potash can also accumulate and eventually foul the tube side of the waste-heat boiler. It is important to have regular E4 maintenance programs to ensure the boiler is clean and foulant free.

High alumina and silica contents in the foulant material indicate that foulant contains refractory fines. The plant has confirmed some refractory damage on the tube sheet. High Fe2O3 content infers that foulant contains corrosion product from the reformer tubes and transfer lines. Ni detected in the foulant sample can come from the corrosion product, since high-temperature piping contains a significant amount of Ni, some might be contributed from the catalyst fines.

After cleaning the tube-side fouling by brushing, the E4 outlet process temperature dropped from about 365°C to 350°C. To prevent waste-heat boiler fouling in the future, the following measures have been planned: • Implement regular waste-heat boiler maintenance programs that ensure the boiler is clean and foulant free.

• Use steam-reforming catalysts containing less and tightly bonded potash if potash fouling recurs after implementing the boiler maintenance programs.

• Add a layer of chloride adsorbent between the hydrogenation catalyst and the ZnO adsorbent beds to prevent corrosion.

• Use a dual-refractory system (insulation layer of silica-alumina and hot-face layer of pure alumina) to insulate the wasteheat boiler instead of shrouded refractory.

• Have adequate flow of blowdown in the waste-heat boiler to prevent shell-side fouling and to ensure good steam quality. HP

Gene Yeh is a registered professional engineer in the state of Louisiana. He holds a BS degree in chemical engineering and ME and PhD degrees in chemical and fuels engineering. He has more than 22 years’ experience in oil refining, catalyst manufacturing and the R&D environment. Dr. Yeh is currently working as an engineering specialist in the process and control systems department of Saudi Aramco in Dhahran, Saudi Arabia. His support areas include catalyst and adsorbent selection, hydrogen plant, hydroprocessing and naphtha reforming.

Said Al-Zahrani is a general supervisor in the process and control systems department of Saudi Aramco with more than 22 years’ rich experience in oil refining. He holds a BS degree in chemical engineering from King Fahd University of Petroleum & Minerals and began his career at Saudi Aramco as a process engineer in the Ras Tanura refinery. Mr Al-Zahrani is the chairman of the multidisciplinary Product Specifications Committee tasked with managing various issues related to Saudi Aramco products and fuel specifications. Mr. Al-Zahrani is a member of several local and international societies, and an officer of American Institute of Chemical Engineers,Saudi Arabian chapter.

Al-Babtain I. Mohammed is a unit supervisor in the Downstream Process Engineering Division of Saudi Aramco. He has 18 years of experience in the refining business and joined P&CSD and participated with FPD and NBD as a technical member in the developing major refining projects and evaluating several technologies.

Nabeel Al-Ghanemi is a process engineer in the manufacturing department at Rabigh Refining & Petrochemical Company since October 2008. He has been involved in troubleshooting hydrogen plant problems in the complex. Mr. Al-Ghanemi holds a B Sc degree in chemical engineering from KAAU, Jeddah, Saudi Arabia.

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FREE Product and Service Information—JUNE 2010 HOW TO USE THE INDEX: The FIRST NUMBER after the company name is the page on which an advertisement appears. The SECOND NUMBER, appearing in parentheses, after the company name, is the READER SERVICE NUMBER. There are several ways readers can obtain information: 1. The quickest way to request information from an advertiser or about an editorial item is to go to www.

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Name ________________________________________________________

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ADVERTISERS in this issue of HYDROCARBON PROCESSING Company Website

Page

RS#

ABV Energy SpA . . . . . . . . . . . . . . .47 (162) www.info.hotims.com/29420-162 www.info.hotims.com/29420-161 www.info.hotims.com/29420-157

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Axens . . . . . . . . . . . . . . . . . . . . . . .88

(53)

Bryan Research & Engineering . . . . .18 (113) C&I Engineering Inc . . . . . . . . . . . . .57 (166) Carver Pump Company . . . . . . . . . .26 (154) www.info.hotims.com/29420-154 www.info.hotims.com/29420-57

Dyna-Therm . . . . . . . . . . . . . . . . . .72 (173) www.info.hotims.com/29420-173

Eaton Filtration . . . . . . . . . . . . . . . . .6 (116) www.info.hotims.com/29420-116

(85)

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(93)

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Paharpur Cooling Towers, Ltd. . . . . .33 (101)

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Johnson Screens Europe . . . . . . . . .27

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Spraying Systems Co . . . . . . . . . . . .31 (71)

HP Marketplace . . . . . . . . . . . . . 82-84

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HCPL . . . . . . . . . . . . . . . . . . . . . . .63 (169) Honeywell Analytics. . . . . . . . . . . . .21

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(79)

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Maxon Corporation . . . . . . . . . . . . .60 (168)

MSA . . . . . . . . . . . . . . . . . . . . . . . .30

KBR . . . . . . . . . . . . . . . . . . . . . . . .12

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RS#

www.info.hotims.com/29420-156

Hunter Buildings . . . . . . . . . . . . . . .59 (167) (76)

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Emirates . . . . . . . . . . . . . . . . . . . . .10

Gulf Publishing Company

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(57)

Page

www.info.hotims.com/29420-168

www.info.hotims.com/29420-169

www.info.hotims.com/29420-166

Company Website

Merichem . . . . . . . . . . . . . . . . . . . .16

Software Reference Showcase . . . .87

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Curtiss-Wright Flow Control Corp . . .2

Gea Wiegand GmbH . . . . . . . . . . . .42 (158)

Software . . . . . . . . . . . . . . . . . . . .81 (177)

www.info.hotims.com/29420-53

Costacurta SpA Vico . . . . . . . . . . . .21

RS#

Construction Boxscore . . . . . . . . . .32 (156)

Aggreko . . . . . . . . . . . . . . . . . . . . .41 (157)

ASCO Valve Inc.. . . . . . . . . . . . . . . .34

Page

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ACS Industries Inc. . . . . . . . . . . . . .46 (161)

Altair Strickland. . . . . . . . . . . . . . . .22

Company Website

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Weir Gabbioneta Srl . . . . . . . . . . . .14

(91)

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Weka AG . . . . . . . . . . . . . . . . . . . .44 (159) www.info.hotims.com/29420-159

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I 85

HPIN WATER MANAGEMENT LORAINE A. HUCHLER, CONTRIBUTING EDITOR Huchler@martechsystems.com

What’s happening with the Legionella Standard?—Part 2 The present draft document provides more than performance metrics, it specifically details a single set of mandatory monitoring practices and corrective actions—the “exclusively correct practices and corrective actions,” without options. By insisting on a “one-size-fits-all” approach instead of matching the intensity of monitoring practices and severity regarding corrective action to the risk of an outbreak, the current draft strictly limits the autonomy of the cooling tower owner to manage the risk of legionella outbreaks.

The draft standard. Close examination of the practices and

corrective actions shows that they are identical in the draft document, and the referenced CTI Guideline. These practices lack a referenced body of evidence of their effectiveness (i.e., lack validation) and generally ignored by cooling tower owners in the US. For example, the draft document requires a single, severe protocol of hyper-chlorination for nonconforming operating conditions such as “poor halogen control” or “frequent hydrocarbon process leaks” or cooling systems that use “reclaimed water or (treated) wastewater as makeup.” Only one of these nonconforming operating conditions is a definitive measurement;1 the remaining operating conditions consist of poorly defined measures of correlated parameters.

In other words, the owner of the cooling tower must act on data that indirectly indicates a possible risk of legionella bacteria in the cooling water and a suspicion that the cooling tower will broadcast these bacteria in the plume and infect persons in the drift zone.

For cooling towers with either high legionella populations, high bacteria populations within 24 hours of hyper-chlorination and “if cases of Legionnaires disease2 are known or suspected and may be associated with the cooling tower,” the procedure requires a single, more severe emergency disinfection procedure. Again, the owner of the cooling tower must act, even if there is a poorly defined suspicion of Legionnaires disease.

It seems prudent to expect the standard to “require proof ” or validation of the risk of potential amplification, broadcast and Legionellosis infection, as well as the efficacy of these high-halogen corrective action treatments: unnecessary corrective actions has consequences. They can cause irreversible corrosion damage to all heat exchangers (carbon steel and especially copper-alloys) and premature replacement of all heat exchangers. This lack of owner control and lack of validation of the corrective actions gives the standard a very punitive tone, with a central focus on high-risk, poorly substantiated “corrective” actions. Standards should not “scare” cooling tower owners into compliance.

1 Legionella

test results greater than 0.0 CFU/ml

2 Legionellosis 3 “Consensus

on Operating Practices for the Control of Feedwater and Boiler Water Chemistry

in Modern Industrial Boilers,” CRTD-Vol. 34, ASME, 1994.

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I JUNE 2010 HYDROCARBON PROCESSING

The greatest challenge for this draft standard may be its design as a highly prescriptive document instead of as an outcome-based document. Prescribing outcomes and allowing options for corrective actions are essential, because the ultimate responsibility for managing the risk belongs to the cooling tower owners.

Consider a different concept. Consider the effectiveness of the ASME3 guidelines in improving the compliance of boiler owners to industry-recommended safe operating limits. The ASME guideline is a “minimalist” document. It prescribes very few practices; it focuses on measurable outcomes and is widely accepted by the technical end-user community.

The proposed CTI legionella standard could also be a “minimalist document” and describe outcomes. Proactive monitoring methods, maintenance practices and new technologies, such as sophisticated diagnostic tools and dispersion models to improve the confidence of infectious disease investigations, would be separate documents that would provide options for cooling tower owners to monitor cooling water, resolve problems and meet the standard’s definitive, measureable requirement of “nondetectable” legionella bacteria. This framework would allow the standard to retain its relevancy as new developments in test methods, modeling techniques and other technologies improve the speed and confidence in diagnosing the source, predicting the risk and prescribing the optimal corrective action to reduce the risk of infection.

In summary. The most important objective of this standard is to proactively balance the risk of an outbreak with the risk of equipment damage from excessively aggressive corrective action.

However, the current draft CTI legionella draft standard requires cooling tower owners to take extreme and well-defined risks for equipment reliability in an effort to mitigate a poorly defined risk without providing efficacy of this damaging remedial chemical treatment. The severe and arbitrary nature of these requirements risks poor acceptance by the end-user community and the lack of definitive metrics for corrective actions make the draft document vulnerable to misinterpretation and legal challenges.

Once published, there will be no opportunity to reshape the standard. It is imperative that the CTI committee re-commit themselves to the fundamental objective of this standard because, ultimately, they serve the industrial community and society. HP End of series. Part 1, May 2010.

The author is president of MarTech Systems, Inc., an engineering consulting firm that provides technical services to optimize water-related systems (steam, cooling and wastewater) in refineries and petrochemical plants. She holds a BS degree in chemical engineering and is a licensed professional engineer in New Jersey and Maryland. She can be reached at: huchler@martechsystems.com.

Showcase

u p s t re a m / m i d s t re a m / d o w n s t re a m

A Supplement to

&

The following companies are display advertisers in the Spring 2010 edition of the Upstream/Downstream Software Reference Guide. You can access the entire guide online at www.gulfpub.com/gpc/. This edition will also be available at many key industry meetings, trade shows and conferences.

Chemstations is a leading global supplier of process simulation software for the following process industries: oil & gas, petrochemicals, chemicals and fine chemicals, including pharmaceuticals. We currently offer several individually licensed, and tightly integrated, technologies to address the needs of the chemical engineer, whether doing new process design or working in the plant.

Heat Transfer Research Inc. (HTRI) is an international consortium founded in 1962 that conducts industrially relevant research and provides software tools for design, rating and simulation of process heat transfer equipment. HTRI also produces a wide range of technical publications and provides other services including contract research, software development, consulting and training. www.info.hotims.com/30874-411

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Codeware is the leading provider of software for the design and analysis of ASME Section VIII pressure vessels and heat exchangers. Since 1985, Codeware’s Engineers have focused exclusively on meeting the engineering software needs of designers and users of vessels and exchangers. Over 1,000 companies currently rely on Codeware products.

m:pro IT Consult is a project services and software products company which enables petroleum refining, petrochemical and other industries to achieve total integration of information sources and applications, from business systems, ERP and supply chain management through to plant information, production planning, scheduling and operations decision support. www.info.hotims.com/30874-402

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The Equity Engineering Group, Inc. is a recognized leader on aging infrastructure fixed equipment service and support for the oil and gas industry. Equity helps plants manage risk and improve profitability with cutting-edge software and consulting strategies that maximize equipment operational availability, control inspection costs and avoid costly shutdowns.

M3 Technoloy is a leading provider of advanced asset scheduling and optimization solutions for oil refining, petrochemical, natural gas-LNG and terminal operating industries. M3’s SIMTO™ software captures eco-nomic opportunities and reduces the cost of managing complex facilities at the plant level, regional operating level and global enterprise level. www.info.hotims.com/30874-414

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geoLOGIC systems ltd. provides well data and integrated software solutions to the energy and production industry.

The company is an innovator in supplying data in more accessible and usable forms so clients can make better decisions - from the well head to senior levels of accounting and administration.

Merrick Systems provides integrated software solutions for oil and gas production operations. Since 1989, Merrick’s product suite has grown to include Best-of-Class applications for field data capture, production hydrocarbon accounting, plant allocations, regulatory reporting, production data management and production monitoring.

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Gulf Publishing Company’s Software Division publishes and distributes more than 40 desktop applications designed specifically for the needs of the engineering community involved in the petroleum industry.

Haverly Systems Inc. is an independent software company specializing in developing optimization-related products and services for over four decades. Their systems are used in over 50 countries by international and independent oil companies as well as petrochemical companies. The effectiveness of their work has long been recognized in the continued patronage and goodwill of our clients. www.info.hotims.com/30874-413

Yokogawa Corporation of America is the North America unit of US $4 billion Yokogawa Electric Corporation, a global leader in the manufacture and supply of instrumentation, process control, and automation solutions. Headquaretered in Newnan, GA., Yokogawa Corporation of America services a diverse customer base with market-leading products including analyzers, flow meters, transmitters, controllers, recorders, data acquistion products, meters, instrumented systems, distributed control systems and more. www.info.hotims.com/30874-407

Improve your swing and maximize your return AxSorb™ is a complete range of high quality activated alumina and molecular sieves These adsorbents have been designed for drying, purification and speciality applications in the refining, petrochemicals and gas processing industries. Squeeze the most from your swing adsorption units and reduce operating costs with AxSorb.

Single source ISO 9001 technology and service provider www.axens.net Beijing +86 10 85 27 57 53 Houston +1 713 840 11 33 Moscow New Delhi +91 11 43399000 Paris +33 1 47 14 25 14 Tokyo

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