HP_2009_06

JUNE 2009

HPIMPACT

SPECIALREPORT

TECHNOLOGY

Mergers, acquisitions review and outlook

PROCESS AND PLANT OPTIMIZATION

Designing stabilization plant filtration systems

Valve demand forecast to turn around

Control, separation, expansion projects

Minimize storage vapor and displacement losses

www.HydrocarbonProcessing.com

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JUNE 2009 • VOL. 88 NO. 6 www.HydrocarbonProcessing.com

SPECIAL REPORT: PROCESS AND PLANT OPTIMIZATION

33 45

Optimize plant performance using dynamic simulation This plant case history illustrates the benefits A. Al-Dossary, M. Al-Juaid, C. Brusamolino, R. Meloni, V. Mertzanis and V. I. Harismiadis

Re-evaluate your expansion projects for the new global market Tighter economic conditions require refiners to reconsider benchmark and optimization tools for revamps B. Fairleigh, J. Jacobs and R. Ohmes

Cover Construction moving forward on the 50,000-bpd coker unit for Total’s $2.2-billion Deep Conversion Project at the Port Arthur, Texas, refinery.

Photo courtesy of Total.

53

Rethink your liquid-liquid separations

61

IT/automation convergence revisited

65

Maximize ethylene gain in acetylene removal units New-generation catalysts and proper operating strategies offer improved selectivity and cycle length M. A. Urbancic, M. Sun, S. Blankenship and D. B. Cooper

19 Valve industry: reason for optimism, despite projected downturn

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Unifying framework for six sigma and process control

21 Common sense needed to forge energy policy, says former Shell CEO

A fresh look investigates general principles in designing process coalescers R. Cusack Keeping automation close-coupled to operation is key A. G. Kern

The advances presented will improve quality and productivity P. B. Deshpande and R. Z. Tantalean

MAINTENANCE/RELIABILITY

79

HPIMPACT 19 Mergers, acquisitions in global oil and gas markets slump

21 Efficiency goals could radically reshape EU’s energy landscape—CERA

Designing and troubleshooting stabilization plant filtration systems—Part 1 Compatability of the filter cartridge media with the plant feed is a major consideration A. Atash Jameh, A. Zamani Gharaghoosh, S. Bazargani, S. Mokhatab and S. Rahimi

GAS PROCESSING DEVELOPMENTS

83

Minimize vaporization and displacement losses from storage containers Consider using this new calculation for recovery A. Bahadori

COLUMNS 11 HPIN RELIABILITY Consider recessedimpeller pumps 13 HPIN EUROPE The future is big; the future is east 15 HPINTEGRATION STRATEGIES Process engineering tools drive efficiency improvements 17 HPIN ASSOCIATIONS Stuck upstream with the downstream blues again

Page 79 Designing and troubleshooting stabilization plant filtration systems—Part 1

DEPARTMENTS 9 HPIN BRIEF • 19 HPIMPACT • 23 HPINNOVATIONS • 27 HPIN CONSTRUCTION • 30 HPI CONSTRUCTION BOXSCORE UPDATE • 86 HPI MARKETPLACE • 89 ADVERTISER INDEX

90 HPIN CONTROL APC designs for minimum maintenance—Part 1

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 Mark Peters mark.peters@gulfpub.com EDITORIAL Editor Les A. Kane Senior Process Editor Stephany Romanow Managing Editor Wendy Weirauch 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 Chris Valdez Artist/Illustrator David Weeks Manager—Advertising Production Cheryl Willis ADVERTISING SALES See Sales Offices page 88.

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HPIN BRIEF WENDY WEIRAUCH, MANAGING EDITOR

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Manufacturers of pumps, compressors and fittings were the largest contingent of exhibitors at ACHEMA 2009, held in mid-May in Frankfurt am Main, Germany. About 950 out of 4,000 exhibitors came from the sector. Solutions that improve efficiency and availability had top billing. In response to rapidly increasing energy costs, more electronic components and sensors are being installed on rotating equipment (pumps, compressors, fans, etc.). “Users are keenly interested in early fault detection systems and predictive systems that provide information on remaining life,” say the conference organizers. For more information on other industry trends showcased at the conference, go to www.achema.de.

Three recent bankruptcies by large ethanol producers have exposed the biofuels sector’s fundamental struggles. These facilities’ financial woes are part of a broader industry downturn that analysts say may claim other casualties. In recent months, the industry has been rocked by volatile corn and oil prices, seen funding dry up for new projects and watched demand ebb as slowing gasoline consumption and lower pump prices reduce the call for blending the fuel with gasoline. About 16% of the total 12.6 billion gallons of corn ethanol production capacity in the US is presently shut down, according to the Renewable Fuels Association. In related news, the US Environmental Protection Agency recently proposed enacting first-ever greenhouse gas performance standards for biofuels.

The standards would take into account all emissions created in the process of making ethanol, rather than just emissions from burning the fuel.

Shipments of industrial controls sank during the first quarter of 2009, according to research from NEMA, a trade association for the electrical manufacturing industry.

That organization’s Primary Industrial Controls Index showed its largest quarter-to-quarter decline on record, contracting more than 23% in the first quarter of 2009 versus the fourth quarter of 2008. Demand for industrial controls has mirrored that of other types of capital equipment over the past few quarters. Inflation-adjusted shipments of industrial control equipment have fallen to their lowest level since 1991. “On a positive note, it appears that the worst of the collapse in manufacturing activity is over,” says NEMA. Nevertheless, demand for industrial controls and similar capital goods will remain weak for the next several quarters, according to the association’s forecast.

■ Demand swings vex refiners The refining industry is grappling with shifting scenarios for tomorrow’s energy landscape. One scenario has the US becoming a net exporter of gasoline by 2010, concludes a new analysis by global management consulting firm Booz & Company.

The report on refining trends explores rising demand for fuel in Asia and the BRIC nations (Brazil, Russia, India and China), mandates for biofuels and the impact of inexpensive automobiles.

“This confluence of factors is confounding an industry that counts on 20-yr predictions to guide investment decisions made today,” according to the report.

On the demand side, the debut of India’s Tata Motors’ $2,500 car has generated a lot of buzz that demand for transportation fuels would increase.

Economic growth in developing economies is another key driver of demand.

In fact, the analysts calculate that just a 1%/yr growth rate among the BRIC countries would add 3 million barrels/ day (MMbpd) of demand for ground transportation fuels by 2025.

An improved performance from banks by the end of 2009 should provide the basis for a pick-up in financial lending. So says a mid-year economic forecast from the European Commission. GDP is projected to fall by about 3% in the US and by a stark 5¼% in Japan in 2009. Although China seems to be in a relatively good position to counter the global recession, growth is expected to slow sharply this year: to about 6%.

Among the five largest EU economies, real GDP is expected to contract in 2009 by about 5.5% in Germany, 4% to 4.5% in Italy and the UK, and by about 3% in France and Spain.

“Looking ahead, still tight financial conditions and weak confidence are set to continue to weigh on economic activity, but they are likely to be gradually offset by the impact of strongly expansionary macroeconomic policies,” according to this analysis.

Energy professionals gathered at largest offshore industry event, held May 4–7 in Houston. Attendance at the 2009 Offshore Technology Conference reached 66,820 participants despite a global economic recession and initial concerns about swine flu. Last year’s exhibition and meetings drew approximately 73,000. One topical luncheon featured Matt Simmons, head of the energy banking firm Simmons and Company International. He discussed his “peak oil” theory—that global crude oil production peaked in May 2005—and noted that US gasoline consumption is growing, and that Korea and much of the rest of the developing world is experiencing strong petroleum demand increases. However, a major industry challenge is the “rust crisis” in the global energy infrastructure. “Leaks are setting new peaks,” according to Mr. Simmons. HP

With a situation of potential oversupply, the possibility looms of the US becoming a net exporter of gasoline and the destruction of refiner margins in developed countries due to the costs of transporting the fuel to BRIC nations, where the demand will be. “Refiners face difficult choices ahead, whether it’s pulling the plug on projects in developed countries, getting into biofuels or expanding into Asia,” said Pedro Caruso, Booz & Company principal.

The report also finds that the pace of capacity addition is picking up, despite high capital costs. Distillation capacity has expanded 3−4 MMbpd over the past four years, and is set to grow approximately 6 MMbpd between 2008 and 2012. HP HYDROCARBON PROCESSING JUNE 2009

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HPIN RELIABILITY HEINZ P. BLOCH, RELIABILITY/EQUIPMENT EDITOR HB@HydrocarbonProcessing.com

Consider recessed-impeller pumps Perhaps you have encountered pumping duties handling freeflowing slurries, sludge and fibrous materials. If used in these services, a standard centrifugal pump may clog, become vaporbound or wear excessively. For the applications mentioned here, pumps with fully recessed impellers (Fig. 1) should be given serious consideration. Recessed “gentle pumping action” impellers incorporate the vortex principle illustrated in Fig. 1 wherein only an estimated 15% of the total fluid throughput makes contact with the fully recessed impeller. These pumps are typically available in flow capacities approaching 100 l/sec (1,580 gpm) and heads ranging to 130 m (430 ft).

Recessed-impeller pumps have been around since the 1930s.

Unfortunately for the user, a number of manufacturers offer recessed-impeller pump configurations that have not advanced from their respective configurational or hydraulic performance constraints for 40 or more years. It is also fair to point out that some legacy models require a degree of maintenance involvement that was considered acceptable decades ago, but is no longer tolerated by best-of-class users.

A number of important characteristics and advancements separate one make or design of recessed-impeller pump from another.

It is worth understanding and considering how recessed-impeller pumps sold worldwide under the Egger or Turo labels have favorably distinguished themselves in this regard. Commercial models became available in the mid-1950s; since then the original Egger design has seen a number of seemingly small, yet important, upgrades. Successive iterations have consistently advanced relevant efficiency and the ability to handle solids with minimum damage to either the pump or the material being pumped. The overall vortex-type operating principle has remained the same.

85%

15%

How recessed-impeller pumps often differ. Most recessed-impeller pumps rotate the liquid and solids inside the casing until the solids reach a speed at which they exit the casing.

This recirculation of solids creates wear in the casing and also increases damage to soft solids. Egger has overcome this problem by designing the casing with an “axial spiral” in the casing. Visualize an automobile tire to represent the basic design of a recessedimpeller casing. Cutting the tire at the top and then twisting it yields a spiral. In like manner, the spiral contour helps guide solids out of the casing; it prevents solids recirculation. The manufacturer has demonstrated on many occasions that this design substantially improves the true overall pump hydraulic efficiency.

Additionally, the axial-spiral twist has greatly reduced component wear and damage to solids being pumped. As a further point of interest, the minimum flow capability of a recessed-impeller pump is much lower than that of conventional radial-spiral casing design pumps. On the minus side, top centerline discharge implies a measure of vulnerability when pumping large hard solids. Solids such as rocks might, on rare occasions, smash through the casing neck. In some rock feed applications, tangential discharge might be viewed as an advantage.

In many cases users and engineering design contractors elect to place emphasis on pump efficiency. When asked to define efficiency, they inevitably refer to power draw. That, unfortunately, is seriously wrong. Some pumps achieve seemingly high hydraulic efficiency by simply letting the impeller edge protrude into the casing. Protruding impellers, of course, limit unimpeded passage of solids through the pump. Reliability professionals are urged to rethink what is of true importance here: the efficiency with which both liquids and solids are being transported. Many “old style” recessed-impeller designs have simply not progressed much since their initial introduction. Their best efficiency points (BEPs) are typically in the range of 30 to 40%. On the other hand, advanced designs incorporating axial-spiral design casing internals and fully recessed impellers will have true and effective BEPs around 50 to 60%. Less energy goes into the liquid and less power is consumed to forward-feed the solids. HP

The author is HP’s Reliability Editor. A practicing

FIG. 1

Operating principle of a radial-discharge fully recessed impeller pump (Source: Emile Egger, Cressier, NE, Switzerland).

consulting engineer with close to 50 years of applicable experience, he advises process plants worldwide on failure analysis, modern lubrication technology, and maintenance cost-avoidance topics. Mr. Bloch has authored or coauthored 17 textbooks on machinery reliability improvement and over 450 papers or articles dealing with related subjects. Most of his books are still in print and are available at www.amazon.com.

HYDROCARBON PROCESSING JUNE 2009

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

The future is big; the future is east Refining margins have been off badly in the first half of 2009.

In Europe, a couple of simpler refiners have, at least temporarily, taken their ball home. Is this just the way Europe weathers a recession, or have the trade winds changed for good?

BP’s Global Indicator Margin is generally taken to be a broad indicator of refinery profitability. When the company announced a 62% drop in net profits for the first quarter, it broke out the downstream margins to illustrate the challenges it faces. For the second quarter of 2008, and the rest of that year, BP pegged northwest European refining margins at $7.35/bbl. By first quarter 2009, that number was down to $4.67. When a disastrous April month is included, the average for 2009 up to April 23 comes out to only $3.07/bbl. The company responded by cutting runs at a number of the 17 sites it operates.

Facility shutterings. Incidentally, BP was not alone in hav-

ing a rocky first quarter. There is also, of course, the price of crude to take into account. BP’s Tony Hayward told the Financial Times that the industry was “adjusting from a $100 oil world to a $50 oil world.”

But what BP’s global indicator margin says about the average refinery hides a darker reality for some of Europe’s aging fleet. In April, Spain’s Repsol announced it would simply stop producing fuel at its Cartagena refinery until margins improve. Across its northern border, Total carried out a major reduction in throughputs as traders began talking of an effort to shut in 15% to 20% of Total’s downstream output in France.

The independent sector in Europe is faring no better. After steadily reducing throughput at its Teesside refinery, reports surfaced in April that Petroplus had shut that site down for the time being. Both Teesside and Cartagena have been the subject of ambitious upgrading plans, but in the heat of the current fire, neither of the simple sites makes very much operational sense.

“The company has flagged up many times that it’s been reducing production rates at Teesside,” an analyst commented. “Margins for refineries like these have been at a dollar or less, which clearly isn’t enough to cover their costs.” The company didn’t return calls attempting to confirm the current situation at the site.

Medium term. A halving in margins is bad enough, but speaking at this year’s International Petroleum Week in London, PFC Energy’s Marc Seris raised an ominous prospect. “We don’t really see a major uptick in demand in Europe after this recession is over,” he warned. “Refineries that shut up shop to get through the downturn are shutting up for good.”

Mr. Seris says that there is a massive migration of refining taking place—from mature markets in the OECD where peak product demand is four years behind us, to places such as India.

Why pin your hopes on the 1 in 10 American adults who don’t yet own a car, when you could bet on India, where only 1 in 10 does?

It’s to that country’s western coast that the gaze of commercial managers and traders is magnetically drawn this summer.

Hefty new supply. For it’s there, at Jamnagar in the Indian

state of Gujarat, that people think big about refining. Once they made brass goods here. Today, there’s the Vadinar refinery, operated by Essar, which at 14 million metric tons/year (MMtpy) is significantly larger than the last refinery built in Europe. But that site is set to grow to a capacity of 34 MMtpy as the owners spend about $6 billion on the site.

But Vadinar is dwarfed by its big brother, the Jamnagar I refinery, operated by Reliance Industries Ltd. At 33 MMpty, Jamnagar I is some 10 times the size of some European refineries.

As you’re probably aware, India’s most valuable company chose not to rest on its laurels and greeted the completion of the Jamnagar site by announcing the construction of its twin, Jamnagar II. The latter has been in commissioning since the beginning of the year, and intends to dispatch exports to as wide a range of deepwater ports as it can. The strategy is to not completely destroy prices at each trading point. With the addition of the new 580,000-bpd/29-MMtpy refinery, a bay in Western India is set to have more refining capacity than the UK. Reliance has established trading offices in London and Houston, and embarked on marketing tours to make its own links with retailing chains and large consumers. The new refinery plans to export a large long-range tanker of distillate every other day.

That’s about 80,000 metric tons—or equivalent to the entire daily throughput of OMV’s Burghausen refinery.

Already the effects are being felt, with Asian refiners feeling the pinch first, as their own sales in the region come under pressure.

But those Euro IV distillates are heading to Europe, while the gasoline is off to the US Eastern Seaboard. That’s two big problems for European refiners right there—the threat to the distillate market, which is Europe’s sole growth market, and the sink for their hapless gasoline surplus.

The only note of consolation for the soon-to-be-contracting European industry is that Jamnagar II couldn’t have come onstream at a worse time for its owners. As a result, Reliance has changed the export-only status of its larger, original refinery to be able to sell fuels into its rapidly growing domestic market.

But in the medium term, the prospects for less profitable US and EU refineries aren’t good. India’s on the march, not to mention the Middle East. HP

The author is HP’s European Editor and 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 2009

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HPINTEGRATION STRATEGIES TOM FISKE, CONTRIBUTING EDITOR tfiske@arcweb.com

Process engineering tools drive efficiency improvements The process engineering tools (PET) market consists of two major segments of software products that play a critical role in designing, creating, commissioning, and operating plants and related assets.

Engineering design applications are used for the physical plant layout; process simulation and optimization applications are used for conceptual process design, process modeling and off-line optimization.

In recent years, PET use expanded rapidly as favorable economic conditions and high demand drove capacity expansion projects around the world. With the onset of the financial crisis in the latter part of 2008, new capital projects began to slow. However, not all industry sectors or regions are affected to the same degree. Many capital projects are still ongoing, so the investments in PET continue. In addition, as the recession reaches global proportions, the market dynamics present owner/operators with new asset life cycle management (ALM) challenges, many of which can be addressed via PET technology.

Laser-scanned data can be processed to generate a photorealistic 3-D model of a plant with accurate dimensional information. Capturing as-built engineering data and CAD modeling for existing facilities reduces risks and enables more projects to be completed on time and within budget. Laser scanning and 3-D plant models have the potential to become standard procedures in brownfield engineering over the next several years.

Laser scanning models are not only useful for brownfield retrofits and revamps, but also for operations. As-built models from laser scanning are finding use within asset management programs.

The models are integrated with other applications for plant maintenance, operations, training, etc., and extend the value of asset data to everyone within the plant.

Support life cycle modeling. EPC and operating companies are not only looking for ways to reduce design and operating costs, but also for ways to maximize and leverage investment in models Focus on collaboration, efficiency and cost savings. developed in the conceptual and design phases of a project. The Major challenges for owners/operators and EPCs include improvmodel evolution approach is an important strategic development that ing engineering and operational performance facilitates model reuse between various simulaby getting the most out of their human and tion and optimization applications throughout physical asset base. To improve engineering ■A model evolution the entire life cycle of a plant or process. A model efficiency, companies are looking for ways to evolution methodology allows models to scale shorten the entire design cycle through concur- methodology allows through the plant life cycle. This includes the rent and collaborative methods. To improve models to scale through research, conceptual design, process design, plant operational performance, companies need design, construction, commissioning, operations more efficient, environmentally friendly and the plant life cycle. and revamp stages. Model evolution maximizes agile assets—along with the design information modeling efforts by reducing duplication and needed to operate and maintain them. ensures that decisions during different life cycle Over the past several years, the trend has been to use every availphases are based on a consistent representation of the plant. able resource to complete projects on time and budget. This means Since detailed engineering models contain a lot of process more partnering projects, both internally and externally. There is a knowledge, companies need to exploit this resource for operagrowing need for smaller companies to assist larger ones and take tional improvements. Operators, supervisors and managers typion some of the engineering design work. Consequently, the project cally run their plants conservatively because they do not want to work is getting more fragmented and distributed—and more chalincrease variability or jeopardize safety. By providing operations lenging to manage. access to a simple-to-use, high-fidelity, model-centric decisionTo better deal with the growing complexity of project work, support system based on rigorous engineering principles, operamany companies are investing in highly sophisticated design, tors can run plants closer to their constraints. HP engineering and collaborative tools to obtain greater efficiency from their valuable resources.

Use and availability of as-built information. There are many risks associated with plant design and construction. The risks are even greater for brownfield engineering projects, where existing structures often clash with proposed changes. Many existing plants do not have the accurate as-built information about their facilities needed to help ensure error-free project execution. With the advent of affordable laser scanning, the process plant engineering industry is quickly embracing the technology as a means to obtain detailed as-built status of existing facilities.

The author is part of the automation consulting team at ARC Advisory Group, Dedham, Massachusetts, covering the global process manufacturing markets. He is responsible for following the global industrial automation markets, writing in-depth research reports, and providing advice to clients, with particular focus on process simulation (including dynamic training simulators), advanced process control, optimization and collaborative production management markets. Dr. Fiske holds a BS from Worcester Polytechnic Institute, an MS from Northeastern University and a PhD from Stevens Institute of Technology, all in chemical engineering. He also holds an MS degree in the management of technology from the Sloan School at MIT and can be reached at e-mail: tfiske@arcweb.com.

HYDROCARBON PROCESSING JUNE 2009

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HPIN ASSOCIATIONS BILLY THINNES, NEWS EDITOR

bt@HydrocarbonProcessing.com

Stuck upstream with the downstream blues again Downstream aficionados need to realize that information pertinent to the petroleum refining business is everywhere, not just within the clearly marked parameters of hydrocarbon processing. For instance, the huge Offshore Technology Conference (OTC) that took place last month in Houston, Texas, is a primarily focused upstream, offshore, E&P gathering. I was privileged enough to have a front row seat for all the proceedings at OTC, given that I was editing the official show newspaper. Since I am a savvy consumer of downstream information, and open to seeing how all facets of the oil business are interrelated, I was able to spy amidst all the pontificating about deepwater drilling and semisubmersibles in the Shtokman Field, some information that refiners would find of interest. I am happy to share it with you here.

Energy challenges. During a panel discussion on the challenges of developing a comprehensive energy policy, two Washington insiders offered analysis.

Jason Grumet, executive director of the National Commission on Energy Policy, discussed the 2008 presidential campaign and subsequent legislative maneuvering in the 111th Congress.

“What was interesting was that you saw a nuanced debate between Obama and Clinton in the Democratic primary,” Mr. Grumet said. “You had the two Democratic candidates disagreeing about the environmental and economic impact of domestic production and it squeezed the debate toward the middle.”

He then went on to analyze what Congress is doing. “It would be nice if we could put together legislation that focuses on both supply and demand, think about our long term needs and have a trajectory where we are reducing carbon emissions while at the same time increasing domestic production in the OCS. There is a space in the debate right now for a comprehensive bill.”

David Holt, executive director for the Consumer Energy Alliance talked about the general public’s current perspective. He

Jason Grumet monitors the intersection of energy issues and politics for the National Commission on Energy Policy.

emphasized that people need to realize that a jump to alternative energy is not going to make an immediate tangible difference.

“We need to grow and increase the supply of alternative energy as a portfolio while we continue to grow oil and gas,” he said.

“Energy is a pervasive issue in society and the more that ambassadors from sessions like this can go out and have an honest discussion with friends and neighbors about where energy comes from and what it means to folks’ daily lives, all that makes a difference,” Mr. Holt said. “All this can’t be solved in Washington. Consumers need to let Washington know that it is time to step up with a comprehensive energy policy.”

Matt Simmons. Everyone’s favorite

“Chicken Little” provocateur in the realm of predicting the decline of energy-related natural resources is Matt Simmons. He spoke during a luncheon at OTC and honed his thesis that the world is now approaching “peak everything.” He took his idea of resources dwindling and applied it to humans, saying that even the shortage in people was getting serious. In regard to current economic woes, he said, “40,000 workers were lost in the in the past 4 months—probably more than all the new hires in 2006−2008. I think that most of these will not return to the oil industry.

Employers expect half of current workers to retire in the next five to ten years.”

In light of the fact that last year at OTC he said that “oil would go through $200 like a hot knife through butter,” it should not be a shock that Mr. Simmons said, “High oil prices were not high enough.”

Mr. Simmons then said that high oil prices failed to dampen demand. He went on to justify his $140-is-a-low-price theme by saying that “prices were not high enough to attract hordes of new hires that were needed.”

When asked to offer solutions to current crises, Mr. Simmons advocated for living in a post-peak oil and gas world. He then enumerated what this means: liberate the work force by eliminating long-distance commuting and let people work at home; grow food locally to substantially limit how much petroleum is involved in food supply; and redesign a new fleet of seafaring vessels to ship people and goods by water.

Climate change. Refiners are stuck in

the middle of the climate change debate, especially with a cap and trade bill beginning to snake its way through US Congressional committees. During a panel discussion on the subject at OTC, Dr. Thomas Peterson, a physical scientist at NOAA’s National Climatic Data Center, said, “The recently observed climate change is beyond the bounds of natural variability.”

Mr. Peterson noted that the global surface temperature time series reveal that the planet has been warming over the last century, and especially over the last few decades.

He noted that most of the globe has warmed from 1900-2008 and that ocean and land temperatures have seen increased warming since about 1980, rising by 1.5° per decade during the past 30 years. “There is a nine out of ten chance that warming is due to humans from 1900–2000,” Mr. Peterson said.

“Global warming is unequivocal,” said Dr. Kevin Trenberth, one of the lead authors of IPCC’s Scientific Assessment of Climate Change in 1995, 2001 and 2007.

“The outlook is for more warming.” HP Perry Fischer and Jerry Greenberg contributed to this column.

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HPIMPACT WENDY WEIRAUCH, MANAGING EDITOR

WW@HydrocarbonProcessing.com

Mergers, acquisitions in global oil and gas markets slump

pared with the final-quarter high of 2006,” according to PwC. Gas grabs top spot. Six of the top ten

Deal activity in the energy sector offered a schizophrenic character in 2008. The first half of the year started with a fairly robust number of transactions, until commodity prices began a steep decline. Then, mergers and acquisitions (M&As) slid alongside them. By the fourth quarter, deal activity had all but dried up, with several potential deals being canceled and no new major deals announced.

“The increase in deal numbers was wholly attributable to upstream activity and smaller deals below $0.5 billion,” according to a report from PricewaterhouseCoopers (PwC).

In contrast, there were significant falls in the number of larger value deals and a big fallingoff of very large deals. Only two deals topped the $5-billion mark in 2008, compared with 10 such deals in 2007 (Table 1).

By regions. Companies continued to

seek growth primarily through acquisition as opposed to exploration as the oil price soared in the first half of the year. Purchases were focused in relatively stable locations such as Australia and Canada as companies looked to safe havens to secure reserves to meet future energy demand.

Like the oil price, 2008 was a tale of ups and downs in deal numbers across different parts of the world. Year-on-year deal numbers were up in all territories with the exception of the dominant North American market and the Russian Federation. Everywhere, though, the pace of deal-making slowed during the year as financial and market conditions deteriorated.

“Companies slammed on the brakes in the final quarter with total O&G deal value down 59% on 2007 levels and 72% com-

2008 were purchases of gas assets. Five of the six were for unconventional resources that require considerable technological investment. All of them were in Australia and North America. This reflects the attraction of targets in stable locations close to end markets as companies responded to security of supply constraints.

The rush to develop Australian coal-bed methane gas assets for LNG export helped catapult Australia’s share of worldwide oil and gas deal value up tenfold. Upstream deal value in Australia multiplied, from $1.7 billion in 2007 to $16.6 billion in 2008.

Outlook. While M&A transactions in the

first half of 2009 look set to remain subdued, it is difficult to see stronger players remaining on the sidelines for the whole of 2009 given the opportunities for acquisitions at low valuations.

PwC expects any easing of debt and equity markets, combined with any positive movement in oil prices, “to herald a reawakening” of deal activity. “The long-term energy supply and demand fundamentals are still compelling. When the market returns, and the financial crisis has passed, the potential for a fast revival in commodity prices and deal-making is there,” says this outlook.

Deals for major assets in locations such as Brazil and Canada, with access to end markets and promising reserve potential, are likely to be high on many companies’ target screens. In the US, the opportunity to access the natural gas shale plays is also very compelling at these valuation levels.

Finally, companies seeking to broaden their portfolios into alternative energy may also find that distress and low valu-

TABLE 1. Analysis of oil and gas deals by sector (2007 vs. 2008 % change in parentheses) Number

Total deal value, $ billion

Average deal value, $ billion

690 (+20%)

114.2 (–14%)

0.166 (–28%)

Midstream

51 (–31%)

12.9 (–57%)

0.252 (–38%)

Downstream

78 (–17%)

17.1 (–72%)

0.220 (–67%)

Services

150 (+1%)

36.2 (–46%)

0.231 (–47%)

Upstream

Source: PricewaterhouseCoopers, O&G Deals 2008 Annual Review

ations in alternative energy stocks offer a ripe opportunity for diversification moves.

“The appetite for moves of this kind will, in part, be influenced by the progress of talks in the run-up to and at the December 2009 UN Climate Summit and the extent to which this establishes a framework for clean energy,” the report concludes.

Valve industry: reason for optimism, despite projected downturn As with many industries, the US and Canadian industrial valve industry saw the steady climb in sales and profits of the last decade turn around in the past year to experience a decline. However, the decrease is slight compared to most other industries, and is not expected to last more than a year.

So says a new report from the Valve Manufacturers Association (www.vma.org).

Valve and actuator shipments rise and fall with the fate of the industries that depend on those shipments. Noting the downward track those industries are now on, the VMA estimates a 5% decrease in industrial valve shipments for 2009—the first such decrease in a decade.

However, 2009 VMA Chairman Sam Bennardo remains optimistic. “In VMA’s market forecast, we estimate shipments of valves and actuators to return to their 2007 levels of $3.8 billion,” he says. “Still, compared to how other industries are faring— the automotive industry, for instance—this is a pretty modest decline, and total sales still should be considerably higher than just five years ago, when total shipments were at $3.2 billion.”

Historical trends. Valve demand nor-

mally does not have the dramatic highs and lows of other industries because the diverse markets that make up end users do not cycle up and down at the same time.

While the global economic downturn is among the worst in recent years, “VMA has more than 70 years of history and statistics that show the valve industry will ultimately thrive because it supplies products for industries that are absolutely essential to a growing domestic and world population,” according to Mr. Bennardo.

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HPIMPACT VMA also predicted that the drop in valve shipments will not continue for more than a year. “The valve industry typically lags about six to nine months behind end users as projects are planned—or canceled,” Mr. Bennardo says. With a huge number of infrastructure projects in the works, and economists predicting a bottoming out of the recession in late 2009, the organization forecasts a possible turnaround in second quarter 2010.

Common sense needed to forge energy policy, says former Shell CEO

example, the planning of a major grassroots refinery, petrochemical facility or liquefaction site for natural gas—from project conception, front-end design, permitting, construction and commissioning—may take 10 years before any salable products are available. These facilities are constructed with an intended service life of 30 to 40 years. Energy policy and planning is very similar to building a new refinery or petrochemical facility, Mr. Hoffmeister said.

Project conditions and economic markets will change as the project moves through the various development phases.

Since the 1970s with the first major oil shock, many failed efforts tried to develop energy policies. The US is living through tough times sourced from poor policies developed and approved in reaction to crisis events.

For example, corn ethanol will be remembered as poor policy to decrease dependence on imported oil. Owners of E85-powered vehicles will be bitterly disappointed after purchasing the first tank of fuel; they may not purchase the second tank of this biofuel upon realizing that they received 25% less mpg with 85% ethanol-blended gasoline, according to Mr. Hoffmeister.

The viewpoint on energy policy is an inverted “bell” shaped curve, according to John Hoffmeister, CEO of the Citizens for Affordable Energy at a recent Energy Hot Topics seminar in Houston, Texas (www. gulfpub.com/events).

The population is split between the two extremes—overzealous environmentalists to the left and die-hard industrialists to the right. Unfortunately for the masses, there is no middle ground on energy issues.

Another important point made by Mr. Hoffmeister, who is the former CEO and president of Shell Oil’s US operations, was that “government should not develop energy policies.” The government has no incentive to educate the public. In the US as well as in other nations, elected government officials are responsible for putting together energy policies and legislation.

Regrettably, these elected officials are disconnected from the gravity stemming from the energy situation and are more closely tuned to achieving their own personal agendas and fulfilling the needs of their constituents. Too often, elected officials work off the next election cycle; they see what is possible by the next election and are acting on very short-term cycles.

The reality is that energy issues and policies are very long-term projects. For

What is the answer? Nations need to look at energy solutions and planning on several levels: short-term, medium-term and long-term. The solution to the complex problem of energy will be accomplished through all three stages.

Energy policy should not condemn a particular energy source. Nuclear power has been completely overlooked in the US after a partial core meltdown at Three-Mile Island in 1979. Yet, nuclear power is an excellent means to supply electrical power and is a much better option to produce electricity rather than burning natural gas.

Likewise, coal is a part of the future energy mix, and clean-burning coal technologies are available. There is not enough wind and solar power to replace fossil fuels. However, wind and solar sources do add value to the nation’s energy resource pool and should be included.

One option to break the dysfunction with energy issues is to create a Federal Energy Reserve Board very similar in structure to the financial Federal Reserve Board of Governors, Mr. Hoffmeister said. There would be a national as well as regional Energy Reserve Boards. These positions would be appointed by the president and serve for a set time. More important, the appointees would

By category. In 2008, automated valves accounted for the biggest share among valve types ($1.24 billion), followed by ball valves ($725 million), and gate, globe and check valves ($584 million). Valve shipments hit their peak in 2008, with $4.0 billion in sales, up from $3.7 billion in 2007.

Of the 15 markets tracked by the VMA, in 2009, water and wastewater is forecast to account for 18% of valves sold, followed by chemicals (16%), petroleum production (12%), petroleum refining (12%) and power generation (11%).

be experts within their fields. Notably, this board would include expert representations from the energy industry (oil and gas, electrical power, nuclear, etc.), the environment, commercial manufacturing, transportation and consumer groups.

–Stephany Romanow

Efficiency goals could radically reshape EU’s energy landscape—CERA Utility companies across the EU registered all-time peak electric and gas usage during the winter of 2008−2009. Action on the EU goal of reducing energy usage 20% by 2020 through energy efficiency could result in significant energy use reduction by 2020, and could fundamentally reshape Europe’s energy landscape by 2030, concludes a study from Cambridge Energy Research Associates (CERA).

However, reduced energy consumption at the level desired by the European Commission (EC) would not be inexpensive— costing at least €250 billion, according to the study. Major new policy initiatives would also be required, making the implications a critical issue for policymakers and industry at a time of great uncertainty.

“The depth of the recession and the resulting freeze-up of investment combined with lead times will mean hard choices ahead for the EU in pursuit of these goals,” said Doug Howe, the study’s director.

“Our analysis shows that if the EU member states pursue both the renewable energy target and the energy efficiency target in a mandatory fashion, and if energy prices begin to rebound in 2010 and later as we think is quite possible, then total natural gas consumption across the EU could drop 16% by 2020 and 35% by 2030 over 2008 levels,” according to CERA. Electricity consumption would likely remain flat. That means overall energy—electricity plus gas— would sustainably decline in this scenario.

“Eye-opening as these results are, even these would not meet the 2020 goal on energy efficiency set out by the Commission, but fall short by nearly half,” added Dr. Howe.

The report concludes, however, that the marketplace alone will not deliver these kinds of savings. The price of energy is too low in most EU countries, and the priority of energy efficiency purchases is relatively low in most households to induce most homeowners and building owners to make the large investments necessary to achieve these goals. HP HYDROCARBON PROCESSING JUNE 2009

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HPINNOVATIONS SELECTED BY HYDROCARBON PROCESSING EDITORS editorial@gulfpub.com

Nanotechnology improves hydroprocessing reactions Criterion has introduced a new technology platform, CENTERA. It is claimed to provide more active and robust structures for hydroprocessing reactions. Applying a breakthrough in understanding of morphology, the active-site assembly of the new catalyst is based on nanotechnology.

With an initial introduction in ultralow-sulfur diesel (ULSD) processing, significant performance gains with CENTERA CoMo and NiMo catalysts of 25% to 50% in desulfurization activity have been demonstrated. Such operating conditions enable improving refinery profitability by maximizing diesel yield, diesel quality and unit reliability.

Key improvements integral to the technology enable assembling and preserving individual active sites. The technology ensures the optimal transformation of oxidic metal nanoparticle precursors to sulfided, active sites. It locks them in place to ensure that high activity is retained. Active-site assembly is the configuration of promoter atoms, since these are the structures where vital hydrodesulfurization, hydrodenitrification and hydrogenation reactions occur.

The first new CENTERA ULSD products provide measurable improvements of up to 25°F over earlier generation products (Fig. 1). Data are claimed to demonstrate that the new technology enhances performance across the full-range of ULSD conditions experienced in the market, from the low-

pressure CoMo catalyst regime of revamped units to the higher-pressure NiMo catalyst environment of grassroots units that process a high percentage of cracked feedstocks.

With activity improvements, refiners can capitalize on upgrading opportunities for ULSD units. Potential flexibilities are claimed to include the ability to increase run length, process more difficult feedstocks and boost throughput. Additionally, the higher activity can be used to reduce the catalyst volume required for ULSD, thus freeing up reactor space to install other upgrading catalysts that can be used for cetane and cold-flow improvements, density reduction and naphtha conversion.

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Mercury analyzer provides fast throughput Teledyne Leeman Labs has introduced the Hydra II mercury analyzer. It is claimed to be an amalgamation of performance and productivity with low detection limits (1 ppt), fast sample throughput and high capacity. The 270-position autosampler enhances productivity in laboratories involved with mercury analysis.

The analyzer incorporates new software that the company says brings unprecedented performance and productivity to mercury analysis, all within a user-friendly graphical interface. This intuitive software automates the analytical process from method development to final report generation. Help is provided in the software,

700 Arab medium SR diesel 300 psig

US cracked feed blend 1,200 psig 25°F

680 670 660 650

25°F

SOR WABT for 10-ppm sulfur, °F

690

640 630 620 610 600 DC-2531

FIG. 1

DC-2618

A comparison of ULSD catalyst performance.

DN-3531

DN-3630

including scheduled maintenance and audio/visual tutorials.

One useful aspect of the Hydra II is its ability to be upgraded to direct solids analysis. This instrument’s ability to perform high-throughput liquids analysis as well as direct solid sample analysis makes it practical for modern labs. Other advantages include: • Wide dynamic range • Over-range protection • High-capacity autosampler • Custom racks for quality control • Direct analysis option.

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New hydroprocessing catalyst enhances performance Albemarle has released an improved hydroprocessing catalyst system called STAX. The system designs are based on reaction chemistry and process conditions in each part of the hydroprocessing reactor. STAX technology can be used to achieve refinery targets that are claimed to include: • Maximum hydrodesulfurization (HDS) with constrained hydrogen consumption • Maximum cetane uplift up to unithydrogen limit • Maximum cycle length at specific sulfur target • Simultaneous sulfur and volume gain target.

Achieving a better understanding of the chemistry of hydroprocessing inside the reactor is a key in designing a catalyst system for maximum performance. Many independent variables are at work at various points within the reactor, and the chemistry taking place at any point both affects and is affected by these variables.

Consider the case of ULSD. The conceptual reaction zone model shows three zones (Fig. 2). Each of the three zones in this 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.

HYDROCARBON PROCESSING JUNE 2009

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HPINNOVATIONS DYNA-THERM CORPORATION

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example has dramatically different reaction conditions that affect catalyst performance.

In zone 1, the primary reaction occurring is direct route HDS. The rate of desulfurization is fast, and sulfur content drops rapidly. At the same time, hydrogen is consumed, so the ppH2 is reduced and ppH2S increases—the latter creating some inhibition.

As the rate of direct route HDS slows, the feedstock moves into zone 2, where the focus shifts from sulfur removal to nitrogen removal—organic nitrogen being the main inhibitor to increased HDS reaction rate. Zone 2 ends when organic nitrogen has been almost completely removed. As the feedstock enters zone 3, hydrogenation route HDS picks up. Although ppH2 is at its lowest in zone 3, the catalyst is operating in a nitrogen-free environment, where the rate of hydrogenation of aromatics—including the aromatic rings of sterically hindered dibenzothiophene molecules—increases. Finally, the feedstock exits the reactor as ULSD. Rather than an average performance level from a single catalyst across all zones, an optimized catalyst system can be designed using catalysts that perform well in a particular regime. Key to STAX technology is claimed to be a model that can predict reaction conditions at each point in the hydrotreating unit.

The technology has been introduced in several commercial units.

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Quality Steam Matters Dyna-Therm custom designs and fabricates horizontal and vertical steam purification drums for your demanding applications. Exit Steam Qualities up to 99.995%.

Benzonaphthothiophenes Dibenzothiophenes Benzothiophenes Sulfides, Mercaptans, Thiophenes Benzonaphthothiophenes Dibenzothiophenes Benzothiophenes Sulfides, Mercaptans, Thiophenes Benzonaphthothiophenes Dibenzothiophenes

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www.DYNA-THERM.com sales@dyna-therm.com Select 152 at www.HydrocarbonProcessing.com/RS

Benzothiophenes Sulfides, Mercaptans, Thiophenes

FIG. 2

Changes in sulfur content and compound types at three reactor stages.

Website focuses on strategies to lower costs, boost production Honeywell has launched TheOptimizedPlant.com, a Website that is claimed to deliver ideas, information and tools to help manufacturers maximize plant performance and get the most out of existing assets.

The site focuses on four key strategies: reducing maintenance costs, reducing risk and improving cash flow, implementing high-ROI solutions and driving down operational costs. It includes a variety of tools including videos, podcasts, white papers, case studies and informational web seminars, all of which offer practical advice for deriving faster returns during lean times.

The site features tips on extending the life of current assets, improving product quality, using certified recycled parts to reduce maintenance costs, installing applications to reduce energy consumption, implementing cost-effective migration strategies, reducing raw material costs, optimizing existing advanced applications and complying with regulations.

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Catalyst promotes diesel make from cycle oil BASF Catalysts is offering a new catalyst technology that is claimed to enable refiners to utilize their current gasoline-oriented fluid catalytic cracking (FCC) units to increase diesel yield. The new catalyst, HDXtra, raises diesel yields by maximizing light-cycle oil (LCO) production from the FCC unit.

This technology, combined with optimized operating conditions, is claimed to enable increased LCO yield up to 10% vol.

Approximately half of the benefit is attributed to the catalyst’s selectivity.

A new global demand for diesel is driving a shift in processing to increase diesel yield from FCC units.

HDXtra offers higher functionality by providing high-matrix activity combined with good coke selectivity. The catalyst also uses moderate zeolite activity. This better controls the amount of LCO cracking into gasoline while also offering low hydrogentransfer activity. Such actions preserve more hydrogen in the LCOs for minimal cetane penalty—a key measure of diesel quality.

New selective matrix cracking technology with the catalyst has been demonstrated in two North American refineries in 2008.

Additional trials are ongoing in Europe and other markets.

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)DEAS FOR GROWTH )DEAS FOR SUSTAINABILITY

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

North America Jacobs Engineering Group Inc. has a contract to provide engineering services for Suncor’s Upgrader 2 (U2) 2010 reliability projects near Fort McMurray, Alberta, Canada. Jacobs Engineering will provide engineering services for a number of reliability projects in Suncor’s U2 heavy oil upgrader complex. The various projects are planned to be completed during the U2 2010 plant outage. Jacobs’ scope encompasses design basis memorandum through detail engineering phases.

Air Products has an agreement to construct a hydrogen production facility in Detroit, Michigan, for Marathon Petroleum Co., LLC. The steam methane reformer will have a hydrogen production capacity of more than 50 million standard cfd and is projected to be completed in the second quarter of 2012. The facility will provide hydrogen and steam for Marathon’s heavy oil upgrade project at its Detroit refinery.

Alfa Laval has an order for membranes and heat exchangers for Novozymes’ new production plant in Nebraska. The order value is about SEK 50 million and delivery is scheduled for 2010. The membranes and heat exchangers will be used to concentrate enzymes for further use in production of starch-based ethanol. The facility in Nebraska will not only supply enzymes for the existing bioethanol industry, it will also produce enzymes to be used in ethanol production from cellulose.

Linde North America and Shell Oil Co. have a long-term contract for Linde to supply hydrogen for Shell’s complex in Deer Park, Texas. The hydrogen will be used primarily for the production of clean-burning transportation fuels at Deer Park Refining Co., a joint venture of Shell and PMI Norteamerica S.A. de C. V. Linde will invest in excess of $50 million to supply hydrogen through a pipeline network connected to its facilities in the Houston Ship Channel. The supply will commence in mid-2010.

South America Invensys Process Systems (IPS) has signed a five-year, $50 million contract to

provide comprehensive safety services and solutions to Petrobras. Under the terms of this contract, the company will implement its safety and critical controls and asset management technology to upgrade and modernize 11 Petrobras refining facilities throughout Brazil. Along with its proprietary technology, IPS will also supply engineering, systems integration and consulting services to help Petrobras meet its strategic goals.

Europe Burckhardt Compression has an order from OJSC Syzran to deliver four process gas compressors for OJSC Syzran’s refinery upgrade project in the Samara region of Russia. The contractor for the project is PMP Ltd. The compressors will be used for a new isomerization unit.

The contract contains two hydrogen make-up process gas compressors and two hydrogen recycle process gas compressors.

The compressors will be skid-mounted by Burckhardt Compression. Delivery will take place in the first quarter 2010, with operation scheduled by mid-2010.

TREND ANALYSIS FORECASTING Hydrocarbon Processing maintains an extensive database of historical HPI project information. Current project activity is published three times a year in the HPI Construction Boxscore. When a project is completed, it is removed from current listings and retained in a database. The 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.

Hydromotive GmbH plans to build a demonstration plant which will produce hydrogen from glycerine at the chemical site in Leuna, Germany. The plant will come onstream in mid-2010 and will reprocess, pyrolyse and reform raw glycerine to produce a hydrogen-rich gas, which will be fed into the existing Leuna II hydrogen plant for the purification and liquefaction of the hydrogen. The liquefied hydrogen will initially be used in German cities such as Berlin and Hamburg, where hydrogen is being used as a fuel.

Middle East Jacobs ZATE has a general engineering services (GES) contract with Rabigh Refining and Petrochemical Co. (Petro Rabigh) for its $15 billion facility in Saudi Arabia. Officials estimate the potential value of the one-year contract at $1.5 million. Petro Rabigh is located in the Red Sea port of Rabigh, on the west coast of Saudi Arabia. The petrochemical complex can produce 18.4 million tpy of petroleum-based products and 2.4 million tpy of ethylene and propylene-based derivatives. Jacobs ZATE will provide mechanical, electrical, civil, chemical engineering for refinery and petrochemical, drafting-related services, management-related services and any other engineering services required by the facility.

Neste Jacobs has a long-term framework agreement with Abu Dhabi National Chemicals Co. (Chemaweyaat) covering engineering services for construction of an industrial chemicals city. The first complex to be built in the Chemaweyaat Industrial City (Madeenat Chemaweyaat) in the Khalifa Industrial Zone at Taweelah in Abu Dhabi, will include an ethylene cracker, a reformer and various polymer and chemicals processing units. Chemaweyaat is planning to build a number of similar complexes in Madeenat Chemaweyaat. Under the framework agreement, Neste Jacobs could be commissioned to carry out significant front-end engineering work.

The Saudi Basic Industries Corp.

(SABIC) and the Saudi International Petrochemical Co. (Sipchem) have signed a memorandum of understanding (MOU).

HYDROCARBON PROCESSING JUNE 2009

I 27

HPIN CONSTRUCTION Under the MOU, SABIC will implement several new petrochemical projects in Saudi Arabia at a preliminary estimated value $3.2 billion, including seven plants for the production of 250,000 tpy of methyl metha acrylate (MMA); 30,000 tpy of polymethyl methaacrylate (PMMA); 200,000 tpy of acrylonitrile; 50,000 tpy of polyacrylonitrile; 50,000 tpy of polyacetyl resins; 3,000 tpy of carbon fiber; and 40,000 tpy of sodium cyanide. Sipchem will build two plants at a preliminary estimated cost of $810 million for the production of 125,000 tpy of polyvinyl acetate and 200,000 tpy of ethylene vinyl acetate. These plants are expected to go onstream by mid-2013.

A SABIC manufacturing affiliate will crack the feedstock allocated to Sipchem and also provide it with ethylene. One of Sipchem’s manufacturing companies will supply carbon monoxide to SABIC for the production of MMA. A new catalytic cracker complex with the capacity of 45,000 bpd is under construction at Abadan oil refinery in Khuzestan, Iran. The project is being financed using a

foreign currency allocation of €350 million.

The new catalytic cracker complex is using UOP technology. The project is expected to become operational in late August.

Asia-Pacific Jacobs Engineering Group Inc. has a contract from Hindustan Petroleum Corp.

Ltd. (HPCL) to provide project management consultancy services for a diesel hydrotreater (DHDT) project at HPCL’s refining complex in Mumbai, India. The estimated overall total installed cost is $650 million. Jacobs will perform front-end engineering design and supervise the lumpsum turnkey contracting for the DHDT project. Work includes the installation of a DHDT and associated facilities in HPCL’s Mumbai refinery to meet Euro-IV specifications for diesel.

Linc Energy Ltd. recently opened its Chinchilla demonstration facility in Australia. The facility features underground coal gasification (UCG) to gas-to-liquids (GTL) technology. It has an UCG gas field, a Fischer-Tropsch GTL plant and an onsite laboratory.

Select 153 at www.HydrocarbonProcessing.com/RS

Kuwait and China have an agreement for a $9 billion refinery project in Zhanjiang, a coastal town in China’s Guangdong Province. The companies involved are Kuwait Petroleum Corp. and Sinopec Corp. The deal will help Kuwait achieve its crude oil export target of 500,000 bpd to China by 2015. The refinery will be designed to process Kuwaiti crude supplied

by Kuwait Petroleum Corp., with a capacity of 300,000 bpd. The ethylene cracker unit is expected to produce 1 million tpy.

Africa The Shaw Group Inc.’s Energy & Chemicals Group has been selected by Sasol to provide basic engineering support for two Fischer-Tropsch units. Shaw is providing the basic engineering package for the Fischer-Tropsch unit of Sasol’s wax expansion project in Sasolburg, South Africa. Shaw’s scope of work includes proprietary equipment design and technology optimization. The project is expected to increase Sasol’s production of medium waxes and liquid paraffins by approximately 50%. HP

Select 87 at www.HydrocarbonProcessing.com/RS

HPI CONSTRUCTION BOXSCORE UPDATE Company

Plant Site

Project

Capacity Est. Cost Status Licensor

Engineering

Constructor

UNITED STATES Alaska California California California Mississippi New Mexico Ohio Oklahoma Oklahoma Pennsylvania Texas Texas

Denali Valero Refining Co Shell Intl Prod Paramount Petr Corp Enerkem Technologies Western Refining Husky Energy Inc Terra Industries Inc Wynnewood Rfg Co ConocoPhillips Flint Hills Resources Ivanhoe Energy

North Slope Benicia Martinez Paramount Pontotoc Gallup Lima Woodward Wynnewood Trainer Caldwell San Antonio

Gas Treating Scrubber Crude Unit Hydrocracker Biofuel Plant Scrubber Hydrocracker Urea Hydrotreat, Gasoline Alkylation, HF Terminal, Petroleum Processing, Heavy Oil

None 75 Mbpd None 25 Mbpd 20 MMgpy 11 Mbpd RE 30 Mbpd 480 m-tpd 13 Mbpd RE 14 Mbpd None 10 bpd

Athabasca Genesee Sturgeon Lake Burnaby Brandon Eider Rock Eider Rock

Processing, Heavy Oil Gasifier Hydrocracker Biorefinery Ammonia Hydrocracker Hydrotreater

20 500 100 100 EX 1350 141 78

La Plata La Plata La Plata Araucaria Linhares Linhares Maua Burgos Cangrejera Cangrejera Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan

Coker, Delayed (replace) Fractionator Gas Plant Ammonia Gas Treating Processing, Oil Cumene Cryogenic Gas Plant Paraxylene Styrene Alkylation (1) Alkylation (2) Coker, Delayed Cracker, FCC (2) Distiller, Crude Hydrogen Hydrotreat, Distillate Hydrotreat, Gasoil Hydrotreat, Naphtha Sulfur

43 250

220

E E U H P H U E E E P C

2010 Belco 2009 CLG

2012 2010 2010 2012 2010 2009

Belco CLG UCSA Axens UOP

Fluor Belco S&B CLG Belco CLG UCSA KP Engineering, LP S&B

S&B

UCSA KP Engineering, LP ConocoPhillips

CANADA Alberta Ivanhoe Energy Alberta EPCOR Power L.P. Alberta Fort Hills Energy British Columbia Lignol Manitoba Koch Chemical New Brunswick Irving Oil Ltd New Brunswick Irving Oil Ltd

bpd MW Mbpd Ml/y m-tpd Mbpd Mbpd

10

E E H C S U U

2015 Siemens CLG 2009 2012 ACSA 2013 CLG 2013 CLG

Ivanhoe Energy Siemens CLG

AMEC

ACSA CLG CLG

ACSA

LATIN AMERICA Argentina Argentina Argentina Brazil Brazil Brazil Brazil Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico

Repsol YPF Repsol YPF Repsol YPF Fosfertil-Ultrafertil Petr Brasileiro SA Petrobras Quattor Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos

RE EX BY EX EX

185 m3/hr None None 1290 m-tpd 18 MMm3/d 200 15 Mbpd 200 110 Mm-tpy 40 200 MMcfd 210 Mtpy 100 Mtpy 13.4 Mtpd 13.4 Mtpd 55.8 Mtpd 42 Mtpd 150 Mtpd 48 MMcfd 37 Mtpd 50 Mbpd 7400 tpd 600 tpy

E E E E 2010 U 2009 C C E H U U U U U U U U U U

FW FW FW ACSA

2008 UOP 2009 2011

FW FW FW ACSA

ACSA

Promon

Platume

2009 2009 2009 2009 2009 2009 2009 2009 2009 2009

DIRECT FIRED HEATERS SINGLE SOURCE ENGINEERING AND FABRICATION

1640 S. 101st E. Avenue · Tulsa, OK 74128

Our comprehensive experience in heat transfer technology and related engineering disciplines ensures that we provide the process industry with state-of-the-art designs and manufacturing. Each system is custom designed for your project specific needs. We work closely with you to optimize the interrelation of thermal, mechanical and structural as well as instrumentation and control engineering disciplines. 3-D modeling of all components is done to prove dimensional accuracy for proper field fit up. Our sister companies, Express Metal Fabricators and St. George Steel, perform the fabrication for all North American projects.

Please forward your requests and inquiries to sales@expresstechtulsa.com or call (918) 622-1420.

www.expresstechtulsa.com 30

I JUNE 2009 HYDROCARBON PROCESSING

Select 154 at www.HydrocarbonProcessing.com/RS

HPI CONSTRUCTION BOXSCORE UPDATE Mexico Mexico Mexico Venezuela Venezuela Venezuela Venezuela

Company

Plant Site

Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Pequiven Pequiven Pequiven not disclosed

Morelos Morelos Morelos Jose Jose Puerto Nutrias Undisclosed

Ethylene (2) Ethylene Oxide (2) Polyethylene (2) Urea (1) Urea (2) Urea Gas Compression (2)

Project

BY 300 Mtpy BY 135 Mtpy 300 Mtpy 2200 m-tpd 2200 m-tpd 2200 m-tpd 25 MW

Capacity Est. Cost Status Licensor H U U E E E P

Naftan Refinery Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Sokolovska Uhelna, a.s.

Neste Jacobs Yara Brunsbuettel MOL Hungarian Oil & Gas ConocoPhillips Raffineria di Gela SpA Eni SpA LyondellBasell/SAT&Co LyondellBasell/SAT&Co LyondellBasell/SAT&Co LyondellBasell/SAT&Co LyondellBasell/SAT&Co LyondellBasell/SAT&Co LyondellBasell/SAT&Co Mazeikiu Nafta Gate Terminal BV TCM StatoilHydro Anwil SA Anwil SA Rosneft Kirishinefteorgsyntez Korimos Togliattiazot

Novopolotsk Burgas Burgas Burgas Burgas Burgas Burgas Vresova Porvoo Brunsbuettel Danube Refinery Whitegate Gela Venice Atyrau Atyrau Atyrau Atyrau Atyrau Atyrau Atyrau Juodeikiai Maasvlakte Karsto Kollsness Wloclawek Wloclawek Achinsk Kirishi Moscow Togliatti

Heater, Vacuum Amine Recovery Hydrogen (1) Hydrogen (2) Offsites Sour Water Stripper Utilities (2) Gasifier FCC, flue gas Urea Hydrocracker Sulfur Recovery Hydrogen Hydrocracker Complex Dehydrogenation, Propane Ethane Cracker Gas Separation Polyethylene (1) Polyethylene (2) Polypropylene Hydrocracker LNG Terminal (2) Amine Gas Plant Nitrogen Oxide Reduction Sys Utilities Hydrotreat, Diesel Hydrocracker Alkylation Urea (3)

None None 7500 kg/hr 7500 kg/hr None None None 200 MW None RE 2000 m-tpd 26 Mbpd 10 tpd 120 m-tpd 21 Mbpd None None None None None None None 35 Mbpd 12000 MMm3/y None RE None None None 35 Mbpd 60 Mbpd RE 400 m-tpd RE 2600 m-tpd

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

Engineering

Constructor

2011 2009 2012 Stamicarbon 2012 Stamicarbon 2013 Stamicarbon 2010

Tecnimont Tecnimont Tecnimont Burckhardt Compression

Tecnimont Tecnimont Tecnimont

2009 2012 2012 2012 2012 2012 2012 2008 2011 2010 2012

FW Technip Technip Technip Technip Technip Technip Siemens Belco UCSA CLG

EUROPE Belarus Bulgaria Bulgaria Bulgaria Bulgaria Bulgaria Bulgaria Czech Republic Finland Germany Hungary Ireland Italy Italy Kazakhstan Kazakhstan Kazakhstan Kazakhstan Kazakhstan Kazakhstan Kazakhstan Lithuania Netherlands Norway Norway Poland Poland Russian Federation Russian Federation Russian Federation Russian Federation

2011 2014 2014 2014 2014 2014 2014 2014 2012 2011 2011 2011 2007 2008 2013 2010

Axens Axens Axens Axens Axens Axens Siemens Belco|SGS UCSA CLG Jacobs Nederland BV Haldor Topsøe CLG LyondellBasell|KPI LyondellBasell|KPI LyondellBasell|KPI LyondellBasell|KPI LyondellBasell|KPI LyondellBasell|KPI LyondellBasell|KPI CLG Aker Clean Carbon

Chemeko ILF Consulting Engineering CLG CLG Exelus 2010 UCSA

Techint CLG

UCSA

Techint

CLG Techint|ENTREPOSE Aker Clean Carbon Aker Solutions Chemeko Chemeko CLG CLG

Neftechimproekt SNKP

UCSA

UCSA

ENTREPOSE|Techint

Remwil

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

Process. Performance.

SOLUTIONS. OnQuest is a leader in process plant engineering and combustion technologies for clients in the petroleum and petrochemical industries. Our expertise includes efficient, energy-saving designs for hydrocarbon processing plants, with specialties in ammonia, hydrogen, syngas, LNG, and ethanol plants.

We are also a world leader in direct-fired process heater technology and burner management systems, and have particular expertise in lump-sum, turnkey projects, refurbishments, and revamps.

To learn more, call Randy Kessler at (909) 451-0502.

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www.onquest-inc.com PROCESS & FURNACE TECHNOLOGIES Select 155 at www.HydrocarbonProcessing.com/RS

HYDROCARBON PROCESSING JUNE 2009

I 31

Spray Nozzles

Spray Control

Spray Analysis

Spray Fabrication

Spray Injector Solutions Improve Performance, Extend Service Life and Reduce Maintenance We have dozens of ways to help optimize the performance of your spray injectors, quills and spool pieces. Here are just a few: U Assistance with nozzle selection and injector placement in the gas stream – critical factors to application success U Validation using 3D modeling capabilities and spray testing in our labs based on your operating conditions ensure performance goals are met U Recirculating, air- or liquid-cooled, multiple nozzle designs and more to meet any quality standard or extreme engineering requirement U Retractable, flexible and multi-directional designs are available to minimize maintenance and service interruptions

Learn More at spray.com/injectors Visit our web site for helpful literature on key considerations in spray injector design and guidelines for optimizing performance.

Our solutions include injectors for: U Distillation columns U Regenerator bypass

Computational Fluid Dynamics (CFD) is often used to help fine tune injector performance requirements and placement

U FCCU water wash U Fractionator water wash U Pollution control equipment U Steam quench U And more

In the US and Canada: 1-800-95-SPRAY | 1-630-665-5000 | spray.com | info@spray.com Select 62 at www.HydrocarbonProcessing.com/RS

PROCESS AND PLANT OPTIMIZATION

SPECIALREPORT

Optimize plant performance using dynamic simulation This plant case history illustrates the benefits A. AL-DOSSARY and M. AL-JUAID, Saudi Aramco, Saudi Arabia; C. BRUSAMOLINO and R. MELONI, Snamprogetti, Milan, Italy; V. MERTZANIS and V. I. HARISMIADIS, Hyperion Systems Engineering, Modeling and Simulation, Athens, Greece

D

ynamic simulation is becoming an important tool for checking out and optimizing plant design leading to improved plant performance. In this work, we present the benefits obtained from applying dynamic process modeling to the compression and medium-pressure steam utility section of a grassroots gas–oil separation plant in the Middle East. Plant construction is in progress and the plant startup is planned for the first half of 2009.

As part of the plant design, the engineering company performed a steady-state simulation of the system to check heat and material balances and equipment sizing. Critical areas were also modeled in dynamic mode and a high-fidelity study was conducted. This study was focused on the compressor and steam systems and included a series of safety and operability scenarios to investigate the plant behavior under a wide range of conditions.

The study examined: • Compressor antisurge, power and pressure limiting controls • Special functions, such as event sequence logic (to allow typical operations to be easily repeated) and a built-in cause-andeffect chart (so that the plant logic is accurately reproduced and monitored).

The main benefits from the dynamic simulation study have been the following: • Establishing the need for advanced control schemes to avoid compressor high discharge and low suction pressures • Verifying relief valve setpoints • Determining tuning constants for the most important controllers • Verifying compressor startup and shutdown procedures • Confirming equipment and valve sizing (including under abnormal conditions) • Verifying that the compressor motors are properly rated to achieve startup • Evaluating different shutdown scenarios for the steam system • Improving plant reliability by defining setpoints that maintain safe and reliable operation of the steam system.

Simulation increases the quality level of plant design activities; benefits are also seen during plant commissioning, allowing for a shorter schedule than would have been the case without the study.

This is because possible startup problems have been identified in advance and the necessary recovery actions have already been planned. Further, it ensures reliable and safe operation of the main plant processes.

In the following sections the possibilities of optimizing the basic design of a process plant well in advance of commissioning are explored. The focus is the dynamic process simulation that can be used to help understand the plant behaviour over time and to ensure continuous improvement in both the plant and procedures.

Dynamic and steady-state simulation. A steady-state model is like a “snapshot” of the unit operation. Any change in the plant conditions, such as changing the pressure specification at the top of a column, requires the model to be resolved. After convergence, the steady-state model should predict where the process will settle. On the other hand, at the heart of a dynamic model, one will find a differential equation integrator. This allows information about the behavior of the unit operation over time to be obtained. All variables are “solved” at each time step and at any specific time the process conditions can be monitored. Compared to the steady-state snapshot, dynamic modeling is more like a movie than a single picture.1 Using a steady state or a dynamic model depends on the exact requirements. For process design, a steady-state model of the unit is initially sufficient. However, when the controllability of the unit is in question, or the process response to transients needs to be investigated, a dynamic simulation is usually needed.

The main difference between dynamic and steady-state modeling is the fact that dynamic modeling is “datasheet driven” while steady-state modeling is driven by thermodynamics and process specifications. This of course has an impact on the level of detail required for successful application of each type of simulation. In steady-state modeling, process specifications are used and the focus is on the feasibility of the process. For example, the temperature on the shell side of a heat exchanger can be specified. Based on the specifications of the nearby equipment, the heat transfer coefficient can be calculated and parameterized, i.e., used as a datum when the system conditions change. In dynamic simulation, the heat transfer coefficient of the exchanger should be first estimated from existing data (e.g., the equipment datasheet and heat and material balance or equivalent). The flows are then estimated based on pressure drops and resistances across the heat exchanger. Only at that stage can the temperature at the exchanger exit be calculated.

There are further examples of differences between the two techniques. In steady-state modeling, pumps are rarely modHYDROCARBON PROCESSING JUNE 2009

I 33

SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION

eled explicitly; flow is possible from low to high pressure since it is determined by thermodynamics alone. Distillation column reboilers and condensers are also typically integrated within the first and last column equilibrium stages. However, in dynamic modeling, resistances across valves and piping are important.

Further, pressure/flow boundaries and exact pump/compressor curves are used.

compressors/pumps and their drivers, actuated valves, etc. Some further important points relating to study scope are: • Volumes are vital for an engineering study since they define the system’s capacity and lag times. Piping isometric data are needed to deduce the volume of the pipework and associated resistances.

• All regulatory control needs to be included in a dynamic model. Advanced process control may be included, depending Typical scope for a dynamic simulation. The typical on the area of the system being modeled. Compressor control scope for an engineering study includes all the items that affect the systems are always modeled since they contribute significantly system’s dynamic behavior. This definition covers modeling the to overall plant behavior. Indeed, an investigation of the perforsystem’s physical limitations and its unit operations such as vessels, mance of the compressor controls may well be one of the aims of the study. Shortcuts may be applied and simplified control schemes may be adopted in some cases, providing that checking the control system itself is not critical. The deciding factor in those cases is if the simplified scheme (which is a subset of the behavior of the installed compressor control) manages to keep the system out of surge conditions.

• The system’s emergency shutdown (ESD) logic is required to capture correctly the effects of a trip or a sudden change in the operational conditions. Often, however, parts of the shutdown logic are disabled in exploratory test runs. Allowing the system to develop in this way without shutdown 1967 Nova Pro Street allows a maximum or minimum pressure/ temperature to be found.

• Valve Cv, stroking time, characteristic curves, failure position and manufacturer are important for an engineering Process Maxum study. Non-return and relief valves are always modeled. For antisurge valves, the Do you have flows up to flow as calculated by the model should 9,900 GPM (2,000 m3/hr), heads up to 720 Ft (220 M), match the datasheet flow for the condispeeds up to 3,500 RPM, and tions provided. temperatures up to 500°F (260°C)? Then you • Compressor and driver data, like inerneed Carver Pump Process Maxum Series muscle! tia and torque versus speed curves for electrical motors, etc., are necessary for accurate With an extended range of hydraulic coverage and rugged construction, the Process Maxum Series is ideal for results.

Industrial Process applications. Manufactured in 35 sizes, • Auxiliary systems are not modeled standard materials include WCB, WCB/316SS, 316SS and since they do not affect the system dynamCD4MCu, with others available upon request. A variety of ics. These typically include the compressor options include various types of mechanical seals and lube oil and seal systems, chemical injecbearing lubrication/cooling arrangements, auxiliary tion, manual valves, drains, sample points, protection devices and certified performance testing. etc. Note that some of the above systems Whatever your requirements, let us build the are indeed modeled if the dynamic model muscle you need! is to be used in an operator training simulator. Minor fittings are not modeled, but their resistance is included in the dynamic model.

Creating Value.

• The thermodynamics and the compoCarver Pump Company nent slate for the system in question are usu2415 Park Avenue ally available before an engineering study, Muscatine, IA 52761 and a heat and material balance, obtained 563.263.3410 Fax: 563.262.0510 from a steady-state model calculation, is www.carverpump.com also available. However, some engineering judgment may be required concerning the number of components to be used in a Select 156 at www.HydrocarbonProcessing.com/RS 34

PROCESS AND PLANT OPTIMIZATION dynamic simulation. The CPU cost of thermodynamic calculations is a strong function of the number of components and large models using many components may end up being too slow for productive use.

Significance for an EPC contractor. For EPC contrac-

SPECIALREPORT

Gas processing system study. The gas processing sys-

tem studied is composed of a series of gas–oil separation plants (GOSPs) downstream of the oil well heads. Each GOSP is composed of: • An oil stabilization system to achieve crude oil, water and gas separation. This includes a series of drums and a crude-stabilization column.

• Two motor-driven compressors, with their related equipment (suction/discharge drums, after-coolers, recycle valve and associated control system). The first compressor is a single-stage low-pressure compressor and the second is a two-stage highpressure compressor.

tors, dynamic simulation is an important tool that allows reviewing engineering activities. Equipment size and design limits are usually defined based on normal plant conditions and good engineering practices. With dynamic simulation test runs, equipment size can be checked in real conditions avoiding estimations and generalized rules. It is also possible to monitor the effect of a single trip on the whole plant and consider all possible scenarios, e.g., the repercussions of one trip in the utility area upon the process area and vice versa. In addition, the automation and control philosophies are monitored. Complex control loops and ESD threshold values are verified before the DCS/ESD factory acceptance test.

In the experience of the authors, the main benefit of dynamic simulation is the opportunity to test the operating procedures (startup, and emergency and normal shutdowns) for the main gas compressors before the real plant startup. Considering the tight commissioning schedules and the criticality of this equipment, dynamic modeling allows familiarization with the compressor performance in the process environment many months in advance THE INDUSTRIAL INSULATION STANDARD FOR THE 21 ST CENTURY of the actual plant startup. There is thus HI G H T E M P E R AT U R E E - G L A S S I N S U L AT I O N ample time to identify potential issues and take corrective actions if necessary.

During the detailed engineering phase, SIGNIFICANT ENERGY SAVINGS any possible scenario can be examined UNBREAKABLE (REUSABLE) and verified to improve plant optimizaSHORTER TURNAROUNDS tion. However, the EPC contractor should DESIGN SAVINGS be conscious of the cost of these activities NO CHEMICAL BINDERS VISIT OUR WEBSITE FOR MORE and the specialist profile of the manpower ph NEUTRAL (7.6) INFORMATION AND FREE required. Therefore, it is necessary to limit the modeling scope by focusing on the critPRODUCT SAMPLES 99.7% WATER RESISTANT ical scenarios. It is also possible to identify TEMPERATURE TO 1380°F the area to be modeled by marking-up only NON-COMBUSTIBLE the PFDs in the first instance. P&IDs may LOWEST CHLORIDE CONTENT be marked-up at a later stage. UnnecesEXCELLENT DIMENSIONAL STABILITY sary details should be avoided and minor LOWEST INSTALLED COST units or equipment should be considered as WWW.HITLINUSA .COM battery limits and/or modeled in a simplified manner. This approach saves time in unnecessary data collection and in model S TA N D O N I T tuning and updating.

Limitations could arise for a licensed RAIN ON IT plant; the licensor (for obvious reasons) will never make available all the documenSURF ON IT? tation requested to support proper modeling. In this case dynamic simulations could be provided directly by the process ( ) licensor or a black-box model could be defined for the scope with a high level of simplifications.

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Select 157 at www.HydrocarbonProcessing.com/RS 35

SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION

GOSP 1

Aux. boiler load control

GOSP 2

Auxiliary boiler A

Condensate stripper Gas-oil wells

Stabilization system

LP

PIC

PIC

V3 V2

Auxiliary boiler B HP1

HP2

Gas treating and export

GOSP 3

Auxiliary boiler C HRSG1

Atmos.

V1 PIT

Users - 4 GOSP (normal) - 1 GOSP (future) - 2 Gas trains

GOSP 4

HRSG2 FIG. 1

Simplified process diagram for the gas processing system.

Excess steam condensers

HRSG3 HRSG4 Anti-surge valve

Feed

FIG. 2

Suction drum

FIG. 3

Discharge drum

Simplified process diagram for a compressor system.

• The gas export pipeline and the metering station.

A simplified process diagram of the process indicating the gas flows is shown in Fig. 1. Note that during normal operation, each compressor stage receives gas from two different locations.

For the typical gas compressor system shown in Fig. 2, the main targets for an engineering study are to: • Verify safe operation during an emergency trip and programmed shutdown, and determine the need for a hot- or cold-gas bypass valve • Confirm that the compressor driver is sufficient for a smooth startup • Prove the suitability of the full startup and shutdown procedures • Review the compressor control algorithm’s ability to maintain compressor operation away from the surge region and choke limit during a series of process disturbances.

A series of compressor trips were examined in the study. The startup procedure was optimized and a number of process upsets were tried and their effects on the operations evaluated. In the subsequent sections, we present some of the results.

Steam system study. A HAZOP study on the steam system

identified that sudden changes in the medium-pressure (MP) steam consumption could lead to potentially hazardous operations for the facility personnel and equipment. Fig. 3 shows the main 36

I JUNE 2009 HYDROCARBON PROCESSING

MP steam to letdown Simplified process diagram for the medium-pressure steam system.

MP steam production equipment and control philosophy used at the water injection facilities and utilities plant.

The MP steam generation system shown in Fig. 3, consists of: • Three auxiliary boilers with capacity automatically controlled by the MP steam header pressure. The control strategy ensures proper load balancing of any given number of auxiliary boilers aligned in the MP steam header.

• Four heat recovery steam generators (HRSGs) with fixed capacity for given ambient conditions • The main MP steam header • MP steam process consumers—four GOSPs in normal/ current operation, one GOSP slot for future operation and two gas trains • MP steam to low-pressure (LP) steam let-down equipment • Main safeguarding equipment: ➤ Excess steam condenser trains, able to downgrade MP steam to condensate during over-pressurization scenarios ➤ Overpressure control valves, relieving steam to atmosphere, and opening sequentially in case of over pressurization ➤ Pressure relief valves installed on boiler/HRSG packages ➤ MP steam high pressure alarms.

In the study, the transient high-pressure build-up in the steam network due to the emergency shutdown of one GOSP train was investigated. Furthermore, the consequences on the MP steam system equipment were evaluated in terms of: • MP steam header over-pressurization effects • Possible mechanical damage to the excess steam condensers due to operation outside of design capacity • Ability to control the load of the auxiliary boilers.

Finally, the performance of existing safeguard equipment was evaluated and recommendations on an operational level were given.

Project management. The EPC contractor organized a team with all the competencies necessary with representatives from the process (gas and auxiliary), instrumentation and automation and machinery departments so that the dynamic study is best

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SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION

organized. The process packages and piping departments were also involved as needed. Due to the complexity of the modeling and the know-how required for compressor modeling and control system integration, it was agreed to involve an external company in this activity.

During the kick-off meeting, the critical cases to be run during the dynamic simulation study were analyzed and defined, and the necessary project documentation (e.g., P&ID, PFD, piping arrangement, control philosophy and description and equipment mechanical data sheets) was made available.

It was planned to organize the activity starting from rev. B or C of the P&IDs with engineering activities well under way. Modifications in engineering and material requisition were not expected but would still be manageable for a few months.

In the third month of the study, the dynamic model was tested during the model validation test (MVT) to provide formal acceptance of the model. During the MVT, parameters like flow, temperature, pressure and main composition were checked with the model in steady-state with reference to the project heat and material balance. Discrepancies were discussed and the actual impact of these on the model was considered. For example, discrepancies occurring in minor chemical components usually have no significant impact on model quality.

All flow lines (including those used only during startup/shutdown/alternate running), equipment sizes and critical parameters were checked in the model and preliminary dynamic performances were monitored with qualitative process indicators.

After a few weeks resolving the MVT punch list, preliminary

40

Head, ft

30

20

10

0 0

FIG. 4

5

10 15 Volume flow, acfm

20

25

Compressor maps (head versus inlet volumetric flow plot) for the LP compressor during startup.

results of critical scenario cases were available, and some months after MVT, a document reporting all the cases (preliminary dynamic simulation study) was issued.

Later, the model was updated considering the latest P&IDs (rev. 0, issued for construction), latest controller setpoint or ESD threshold values and the mechanical data sheets coming from the vendors. The issue of piping isometric drawings for the compressor area was organized in ® advance to provide correct volumes for the MICROTHERM dynamic simulation study.

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0.120

0.140

Thermal Conductivity (W/m-K) at 600 °C Mean

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0.160 Data Per ASTM Testing Standards

and HP compressors have a vendor-declared startup pressure lower than the settling-out pressure. Furthermore, the HP compressor has a startup pressure that is significantly lower than the normal operating suction pressure, due to the motor sizing limitation.

For these reasons, detailed startup procedures have been developed that include the following major steps: • Compressor system depressurization • Drainage • Purge • Pressurization • Auxiliary system startup • Main motor startup • Compressor loading.

A detailed dynamic simulation study was performed to: • Verify the proposed sequences and confirm that the compressor systems can start up without problems.

• Review performance of the compressor antisurge system (valve and control tuning).

Select 57 at www.HydrocarbonProcessing.com/RS

SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION 40,000

16

35,000

25,000 Ft

Pressure, psig

12

LP

30,000

20,000 15,000 10,000

8

5,000 0 60,000 155

Time

50,000 40,000

Suction pressure of the first (blue line) and second (pink line) HP compressor stages versus time during startup.

Ft

FIG. 5

105

20,000

400

0.71

30,000

HP1

4 55

10,000

300

200

0.67 LP

0

20,000

HP1

100

Time

200

Ft

15,000

100

0.65

0.63 0

25,000

HP2

HP2

Pressure, psig

Pressure, psig

0.69

10,000

0 300

5,000 0

FIG. 6

Suction pressure of all compressor stages versus time. The LP compressor (blue line) is linked to the left-hand axis, while the HP compressor first and second stages (pink and green lines respectively) are the right-hand axis.

• Observe and analyze transients.

Starting from compressor shutdown conditions, the system was depressurized to the startup pressure value. All critical valves were positioned (opened or closed) based on the startup sequence procedure. Manual steps, like casing draining, were not included.

The motor was then started with the compressor running in full recycle. In this phase, the critical equipment was checked for safe operation with regards to surge, motor overload and subatmospheric conditions at the compressor suction.

After motor startup, the system loading phase was checked. It was found that simply opening the suction and discharge isolation valves was not enough for a smooth startup. The startup procedures have been modified so that: • The HP compressor starts up with the interstage isolation valve closed. This is required to avoid vacuum in the first-stage suction. The interstage valve can be opened when a low differential pressure across that valve is achieved.

40

I JUNE 2009 HYDROCARBON PROCESSING

FIG. 7

Compressor maps (head versus inlet volumetric flow plot) in the case of flow reduction to 30% of the normal operation. The blue arrows show the movement of the operation points.

• Opening the LP compressor suction and discharge isolation valves should happen only when the pressures in the HP compressor section are balanced. In general, the isolation valves have to open in a way to avoid both massive flows and vacuum at the LP/ HP compressor suction.

Thus, the simulation results in this case can be summarized as: • The startup sequences were modified and validated.

• The inlet valves’ sizings were verified and their optimum opening ramps, during the loading phase, were defined.

• The antisurge valve sizing and control were shown to be suitable—there was no surge during startup. A typical startup compressor map is presented in Fig. 4 for the LP compressor. The blue line is the locus of the operating points and the pink line signifies the compressors’ surge and operating line at nominal speed.

PROCESS AND PLANT OPTIMIZATION

FBM HUDSON ITALIANA, established in 1941, is today a worldwide leading brand in manufacturing of process equipments for Oil & Gas and Petrochemical field.

FBM HUDSON ITALIANA, a member of KNM Group since April 2006, has become the Group’s Core Centre for Engineering Excellence in terms of design and technology thanks to its engineering expertise of over 60 years in this business contributing significantly to the growth and development of the products for the Group.

FBM HUDSON ITALIANA is specialised in the research & manufacture of: • Air Cooled Heat Exchangers • Process Gas Boilers • Highly sophisticated S&T Heat Exchangers • High Pressure Urea & Ammonia Exchangers • Welded Plate Heat Exchangers • After Sales Service • Spare Parts The synergy in terms of production, customer base, engineering skills and financial enable the Group to achieve its target to be a One Stop Centre for our clients in supplying the structure and expertise of an international group spread over 16 manufacturing facilities and Engineering offices across the globe in 10 countries granting KNM a very good knowledge of local needs together with matchless know-how.

Pressure, psig

• The possibility of vacuum at the com- installed for the HP compressor. pressor suction during startup was avoided, Thus, the simulation results can be sumas seen in Fig. 5 for the HP compressor. marized as follows: Compressor net flow reduction. Both • The size of the antisurge valve was the LP and HP compressors could be forced sufficient so that any surge during flow to operate at different flowrates in a 30%– reduction was avoided.

100% range of the maximum (contractual • There was a need to include discharge design basis). This wide operating win- pressure limiting control functionality in dow is required to satisfy plant turndown the LP compressor control scheme. requirements. However, the compressors • A compressor shutdown or instability is normally operate in a 60%–100% range not foreseen. This is demonstrated in Fig. 6 of the maximum operating flowrate. This where the compressor stage suction pressure flow range is much smaller than the plant is plotted against time (blue line: LP comturndown requirements. For these reasons, pressor; pink line: HP compressor first stage; a test has been performed to: green line: HP compressor second stage).

• Verify that the compressors can withThe movement of the compressor operstand a significant flow reduction and are ating points to lower flows is clearly seen able to run at turndown conditions in the compressor maps in Fig. 7. The blue • Review the compressor antisurge sys- lines are the loci of the operating points tem performance (valve and control tuning) and the pink lines signify the compressors’ • Observe, analyze and understand the surge and operating lines. transients.

The base dynamic model describes nor- HP compressor limit variation. Gas mal operations. Simple ramps were used from the HP compressor is sent to a manito reduce simultaneously all of the incom- fold and then to the plant gas treatment ing feed flows to 30% of their normal section. A minimum compressor discharge operational values. During this phase of pressure needs to be guaranteed (contracthe test, the results were checked to verify tual design basis). For that reason, the HP safe operation of equipment in terms of compressor system is equipped with a dissurge occurrence, compressor instability or charge pressure control valve. Appropriately compressor trip. designed dynamic simulation tests were Simulation results showed the engage- performed to understand the compresment of the antisurge control system to sor system response in the case of battery avoid surge and the discharge pressure limit pressure variations. The base dynamic increase for both compressors. More spe- model describes normal operations. A simcifically, the LP compressor discharge pres- ple ramp was used to simulate increasing/ sure came quite close to its high-pressure decreasing battery limit pressure. trip limit. The increased possibility of trigThe obtained results showed that: gering compressor emergency shutdown • When the battery limit pressure was dictated the need to test the functionality of the LP compres480 sor system after the addition of a discharge pressure limSP 475 iting controller. This additional controller acts on the antisurge MV valve and limits the 470 discharge pressure.

In this way the pressure fluctuations that 465 could potentially lead to a compressor shutdown can be avoided.

460 Similarly, it was 0 50 100 150 200 verified that a disTime charge pressure limFIG. 8 Compressor system discharge pressure (MV) and its iting controller was setpoint (SP). not required to be

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A Production/Yield Accounting system that becomes the foundation for your loss control initiatives. It assists you with your daily sitewide mass balance on a tankͲbyͲtank and unit level. A discrete events simulator is also available to check the feasibility of your operations schedule including your docks, tank yards, process units and pipelines.

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decreased, the HP 30 discharge pressure control valve closed to maintain the desired compressor Boiler B discharge pressure.

The valve design was 20 20% MCR found to be adequate, while the controller tuning was important to minimize downstream pressure oscil10 lations. This is shown in Fig. 8, where the Boiler C pressure downstream of the HP compressor is plotted against 0 time (dark blue line).

0 100 200 300 400 The green dashed Time line is the controller FIG. 9 Auxiliary boiler load versus time in the case of one GOSP setpoint. The comtrain emergency trip. pressor operating points hardly move at all. results showed that a single auxiliary boiler • When the battery limit pressure was capable of maintaining the MP steam was increased, the overpressure control header pressure control. This is presented valve opened to flare the flow that was graphically in Fig. 9, where the percentage not delivered to the downstream system. of boiler load is plotted against time. The This avoided compressor instability and/ dark blue line represents boiler C load, or shutdown. The overall flow through the while the pink line shows the boiler B compressor system was reduced due to the load. The minimum guaranteed automatic higher head of the HP compressor. control threshold is shown in the dashed green line.

Steam system—one GOSP shutBoiler B is the one that remained in down case. The operational setup that operation. As can be seen from the plot, was investigated reflects the system future boiler B load remained marginally above operation in summer ambient conditions. the minimum guaranteed automatic conIn this case, two out of the three auxiliary trol threshold during this scenario. Evenboilers run at 25% of the maximum con- tually, boiler B load settled quite close to tinuous rating (MCR). All four HRSGs the threshold. These results highlighted the run at their fixed loads and all five GOSPs, need for tight MP steam header pressure along with the two gas trains, are consum- control and subsequent auxiliary boiler ing MP steam. load control.

The HAZOP study identified the At the same time our results showed emergency trip of one GOSP and the that: subsequent over-pressurization of the • The MP steam header was not presMP header as a possible cause for con- surized above its design pressure. cern. However, the dynamic simulation • The capacity and pressure of the revealed that during such a transient, the excess steam condensers did not exceed load for the two operating auxiliary boilers their design limits. would drop below the minimum guaran• The overpressure valves to atmosphere teed automatic control threshold (20% of remained closed and the pressure relief MCR). Eventually, both auxiliary boilers valves were not lifted during this scenario. would be forced to trip due to low fuel gas pressure.

Other considerations and future Therefore, the issue of steam produc- work. Dynamic simulation allowed anation sustainability became critical and lyzing several critical plant scenarios. In six dictated the need to test an alternative or seven months it was possible to review a scenario in which the emergency trip of series of options and provide answers and/ one GOSP was followed by the sched- or alternatives to a number of challenging uled shutdown of one auxiliary boiler. The issues. Overall, this activity improved proBoiler load, %

Efficiency Improvements

PROCESS AND PLANT OPTIMIZATION

PROCESS AND PLANT OPTIMIZATION cess quality and plant knowledge. Due to the early identification of potential challenges, a smoother startup with a shorter schedule is expected.

It is strongly recommended to use and continually improve the dynamic model after plant commissioning based on the real plant performance and feed compositions. Data should be collected using a management information system after the plant has been stabilized. Several operating conditions should be used after a careful validation and reconciliation. It is also suggested to use the dynamic models for a “what if ” analysis during plant control optimization, and also for developing an operator training simulator. However, it is important to review and verify the simplifications and assumptions considered and ensure that the model is suitable for these other purposes. For example, in an operator training simulator, modeling the chemical injection area or spare equipment is required, and clear graphics and user-friendly methods to change boundary conditions are expected.

On the other hand, these are not needed for an engineering study. HP

SPECIALREPORT

Vasileios Mertzanis is a supervisor engineer at Hyperion Systems Engineering, Athens, Greece. He has three years’ experience in the edible oils processing industry and eight years’ experience in the oil & gas industry with particular emphasis on dynamic model development using the most commercially known engineering platforms and on operator training simulators project engineering. Mr. Mertzanis holds an MSc in chemical engineering from the University of Patras, Greece.

Vassilis Harismiadis is a business development manager at Hyperion Systems Engineering, Athens, Greece. He has nine years’ experience in the oil & gas industry with particular emphasis on using dynamic process modeling to improve plant effectiveness.

Dr. Harismiadis holds a PhD degree from NTU Athens in thermodynamic modeling of complex systems.

ACKNOWLEDGMENT Revised and updated from an earlier presentation at the Middle East Petrotech Conference, May 25–29, 2008.

1

LITERATURE CITED Psarrou, S., Bessiris, Y., Phillips, I. and Harismiadis, V. I., “Dynamic simulation useful for reviewing plant control, design,” Oil & Gas Journal, August 13, 2007, volume 105, issue 30.

Abdullah Al-Dossary is a process engineer at Saudi Aramco. He has more than 10 years of experience in the oil & gas industry with particular emphasis in gas oil separation and crude processing. Mr. Al-Dossary holds a BS degree in chemical engineering.

Mazen Al-Juaid is a process engineer at Saudi Aramco. He holds a BS degree in chemical engineering and has six years of experience in engineering and development of industrial utilities plants at Saudi Aramco.

Cristian Brusamolino is a technology manager at Snamprogetti Oil & Gas Treatment Department, Milan, Italy.

He has more than 11 years’ experience in oil & gas industry process activities, with particular emphasis in gas oil separation plants. Mr. Brusamolino received his doctorate degree in chemical engineering from Politecnico of Milan (Italy) in 1994.

Regina Meloni is a senior control engineer specialist in Snamprogetti’s Automation Department, Milan, Italy.

She has more than 20 years’ experience of oil & gas industry automation and dynamic simulation with particular emphasis in the area of operator training simulators. Dr. Meloni received her doctorate degree in chemical engineering from Politecnico of Milan (Italy) in 1983.

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43

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PROCESS AND PLANT OPTIMIZATION

SPECIALREPORT

Re-evaluate your expansion projects for the new global market Tighter economic conditions require refiners to reconsider benchmark and optimization tools for revamps B. FAIRLEIGH, KBC Advanced Technologies, Inc., Parsippany, New Jersey; J. JACOBS and R. OHMES, KBC Advanced Technologies, Inc., Houston, Texas

H

eightened activity levels within the refining industry were driven by high margins, facility expansions and unit revamps, and increasingly more stringent environmental regulations. Events from late 2008 and early 2009 signaled a new phase for this industry. As the global economy slowed, refining margins tightened and capital projects are being reexamined.

The changing political leadership in the US promises more emphasis on curtailing greenhouse gas (GHG) emissions, which will impact daily refinery operations and the final product mix of petroleum-based transportation fuels. Current changes in the financial markets make strategic planning even more problematic. Now refiners must apply a disciplined, proven approach to evaluate various options for profit improvement and environmental compliance.

In the present dynamic market, refiners need a profit-enhancement process that focuses on achieving sustained profitability for both the short- and long-term. The major components of this process include: • Benchmark and set baseline • Identify opportunities • Evaluate low-cost options • Evaluate capital project options • Implement the process • Ensure sustainment and continuous improvement.

Several key success criteria for utilizing this process, with particular attention directed to proper benchmarking, application of tools, creating effective teams and incorporating market impacts will be presented.

Changing market environment. Before examining the

profit-enhancement process, reviewing the current and future state of the refining industry is appropriate. Over the last several years, particularly 2008, the refining industry endured dramatic changes. Since 2000, the expanding global refined product demand (especially from China and India) consumed the global spare refining capacity. The industry shifted from an era of tight margins in the 1980s and mid-1990s to what some would deem a “golden age” of refining. Cracking spreads achieved record levels, and refiners focused on maximum throughput and utilization.

For the first time in many years, refiners were examining expansions of existing facilities and building new refineries to meet anticipated demand.

TABLE 1. Structural and market forces on the global refining industry—past and present Past: 1980s/1990s

Current and future: 2009 and beyond

Crude sources

Largely light and medium conventional crudes

Shifting to heavy/sour crudes, including high TAN and synthetic crudes

Global product demand

Flat to slight growth

Falling due to economic slowdown but long-term projection to increase ~1% per year2

Product demand mix Gasoline >> diesel

Diesel >> gasoline

Markets

Local

Global, with large movements between regions

Biofuels

Limited usage

Significant volumes due to legislative requirements

Refining capacity and utilization

Significant global overcapacity

Relatively tight but projected to increase with expansions

Environmental

Air and water

All sectors, including GHG, carbon emissions, etc.

Key specifications

Low-sulfur diesel and reformulated gasoline in some countries

Global adoption of ultra-lowsulfur gasoline and diesel, oxygenates in gasoline, highcetane diesel, biofuels, lowersulfur fuel oils

Capital projects

Environmental and “stay in business” only

Environmental, expansion, alternate crude slates

Staff demographics

Traditionalist giving way to Generation Xs, “buyers” job market

Baby Boomers giving way to Generation Y, tight job market

Data, tools and computing

Computers becoming part of normal job, relatively scarce data, LPs and simple models

High powered computing readily available, data overload, complex LPs and rigorous full refinery simulations

However, the refining industry is possibly heading for a new dramatic business cycle, which started in late 2008 and is very reminiscent of the 1980s and 1990s, with some stark differences.

High prices for transportation fuels and a weakening global economy have severely curtailed demand and demand growth HYDROCARBON PROCESSING JUNE 2009

I 45

SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION

in the US and Asia-Pacific. Many refiners are facing breakeven margins and are scaling back production levels. Under these conditions, production philosophies are changing from “more, more, more” to “optimize what you have.” Capital investment plans are frozen or being reexamined in light of record high commodity prices, tight capital markets and growing alternative fuel regulations.

Before refiners revert to the strategies and practices from 15 years ago to meet this latest philosophy, some important structural and markets changes, as summarized in Table 1, should be considered. As refiners consider optimizing existing facilities or examining expansion options, the profit enhancement process should account for several issues: • Converting raw data to useable and actionable benchmarks is a critical first step, particularly for the newest generation of engineers.

• Creating effective teams with the proper skills and expertise is essential.

• Applying the right tools at the proper time helps to properly analyze and account for changes in crude and product slate shifts.

• With diesel margins projected to remain strong, examining operational and capital investment opportunities could lead to profitability in the present tight economic environment.

• Properly evaluating capital projects is important under this era of volatility for commodity pricing, engineering costs and future margins.

Road map for success. Virtually all refiners accept the prem-

ise that their refinery needs goals of optimization and continuous improvement for profitability sustainment. What may not be

Sustainment and continuous improvement

Project teams

Implement Evaluate capital projects

Evaluate low-cost options Identify opportunities Benchmark and set baseline FIG. 1

Profit enhancement team, tools and methodologies

The road map for successful for plant projects involves six steps.

agreed upon, particularly within an organization, is the process to be used on achieving these goals. Fig. 1 summarizes a road map that can be successfully used for many industrial facilities. The six main steps that define the profit-enhancement process are: Step 1. Benchmark and set baseline. The first step in any journey is to understand the current situation. Defining the status of the facility via benchmarking becomes necessary. Setting a baseline operation provides a basis of comparison for future operations. Test runs and/or unit monitoring spreadsheets are part of the baselining effort. The baseline refinery operation can be represented in models such as a refinery linear programming (LP) model or a nonlinear simulation flowsheet.

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PROCESS AND PLANT OPTIMIZATION Step 2. Identify opportunities. The next step focuses on identifying opportunities. Having a team with broad organizational and experience representation is the key to completing the brainstorming activity and selecting opportunities with the most promise for evaluation.

Step 3. Evaluate low-cost options. Once the baseline operation is agreed upon and potential opportunities are identified, the next step is to evaluate low-cost (or no-cost) options. This step is really part of an overall refinery optimization using existing assets. Some items may be no-cost (i.e., cut-point changes, higher or lower reactor severity, tower reflux optimization, etc.) or have minimal cost (replace control valve, pipe jumper, etc.) Focus on optimizing the existing facility before examining significant investment is necessary to fairly evaluate capital projects.

Step 4. Evaluate capital projects. When profit enhancement team members are confident that they have defined a mechanism to optimize their existing plant, the team can focus on capital projects. In this stage, the team considers strategic investment opportunities to meet market and environmental needs while dovetailing with existing facility optimization efforts. The optimized base facility serves as the proper basis for evaluating projects, which could include unit revamps and/or refinery reconfigurations.

Step 5. Implement. A fully developed profit improvement plan does not pay off until opportunities are implemented. An implementation plan includes a priority list of items and projects to complete and the formation of project teams charged with implementing these opportunities.

Step 6. Sustainment and continuous improvement. As with any successful process, a feedback loop is needed. For this process, the sustainment and continuous improvement step ensures that the organization examines itself against accepted benchmarks and baselines while looking for future opportunities and accounting for market changes. The key to this step is a tracking program so that staff and management can periodically review organizational progress.

Following this road map (Fig. 1) in the correct order typically provides the best return on the amount of time invested by the profit enhancement team. The temptation is strong for team members to jump directly to the development of opportunities (i.e., Steps 3 or 4) before setting the baseline. Doing so can be a mistake, since opportunities are rarely accepted by management until they are compared with a validated baseline operation.

The easier and less time-consuming path to develop viable and believable opportunities is after the benchmarking and baseline step (Step 1) has been achieved. Optimizing existing assets before evaluating large strategic projects ensures that the value from large

SPECIALREPORT

capital investments is truly justified. Such an endeavor can not be done by a single person; it requires a team focused on optimization and improvement.

Creating an effective team. At the heart of this process is

an effective multi-disciplinary team. Although most refineries are adequately staffed for normal operations and troubleshooting, sufficient resources are not available to develop and carry out major profit improvement initiatives. For example, an effective strategy for project work is to make a clear division of manpower between the “project team” and the “plant operations/maintenance” team.

However, a profit enhancement team will still require input from

Select 177 at www.HydrocarbonProcessing.com/RS

47

SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION

a wide variety of knowledgeable employees, not just those who are selected for a particular project.

The profit enhancement team makeup is 1) a core group guiding and executing the bulk of the activities and 2) support groups providing subject matter experts. Being a member of such a team would not need to be a full-time job, but would require meeting periodically with the team leader or moderators to assist in making improvements and reviewing results. The profit enhancement team would also be empowered to call upon various expertise within the refinery organization as needed.

A proven method is to set up a profit enhancement team that is moderated by consultants and includes refinery employees and possibly staff from licensor(s), catalyst vendor(s) and engineering and construction firm personnel. The key is to strike a balance between using internal resources with knowledge of the facility and process and using external resources bringing a broader view to inject new ideas and to challenge perceived limits and constraints.

the downstream units’ performance—reformer and isomerization units. An increase in the naphtha splitter reflux ultimately increased the volumes of gasoline and benzene, which were the highest value products. The additional benefit from the yield improvement was much greater than the cost of the higher reflux ratio in the splitter.1 3) In cash constrained environments, the temptation is to minimize cost by limiting FCC catalyst additions and lowering equilibrium activity targets. The optimal microactivity testing (MAT) target is dependent on many operating parameters, and lowering activity targets can adversely impact yield selectivity and volume gain.

These examples illustrate the need for operators of a profitfocused refinery to see the “big picture” and to consider valuefocused analysis methods over cost-focused measures. Even experienced operators and engineers may not always be able to see the “knock-on” effects of changing various operating parameters.

When changing one variable affects dozens of others, how can a refiner be certain that the most optimal move is being made?

Several resources are available to aid in this effort and these include: • Unit test runs • Licensors/technology suppliers • Linear programming (LP) models • Unit and refinery simulators.

Unit test runs can be used to test operating conditions and the effects on unit performance. Changes can be observed on the facility’s own processing units. However, test runs may require a significant amount of coordination and planning by engineering, operations and laboratory analysis to be effective.

Licensors and suppliers can also help to understand the variables associated with optimal operation. The technology licensor can estimate the impacts of new feeds or operating strategies on unit performance. However, licensor recommendations are often centered on specific technologies that they provide. Unfortunately, these resources cannot be relied upon for continuous support when optimizing the facility.

Refinery LPs are an excellent tool to understand the systemic interactions within the refinery. An LP works to optimize the refinery profitability within a set of feed, economic, operating and product constraints. It also provides a robust platform to analyze refinery operations with varying feedstocks, economics and operating parameters depending on the structure and complexity. The major downside of an LP is that the linear structure limits the predictive capability for large step outs, as this model relies on linearization of a nonlinear process. Also, proper maintenance of the LP is required to maximize its effectiveness. With the right expertise, these gaps can be mitigated by the interactive use of unit simulators.

‘Fit-for-purpose’ tools and methodologies. Many tools exist to aid in the identification and evaluation of optimization opportunities. Applying the right tool for the proper time helps to analyze and to account for changes in crude and product slate shifts. Many refinery operators attempt to achieve optimization with unit KPI tracking, crude selection with a refinery LP and various cost-reduction programs. These activities are necessary components, but optimization benefits from these methods alone will be limited by the overall complexity of refinery operations and system interactions, coupled with the constantly changing economics and regulatory requirements.

Cost reductions, particularly during lean margin periods, are popular with refiners since these changes are often more readily measurable and visible. However, what about the impact on the overall facility profitability? Consider these examples, which have been witnessed in various refineries: 1) To conserve energy, steam was minimized at a fluid catalytic cracking (FCC) feed injection point. A unit test run demonstrated more valuable yields from the FCC were realized with a higher steam injection rate. The enhanced profitability from the FCC more than compensated for the extra costs in steam. Steam usage as a wt% of feed may be a more effective measure than the steam injection rate to approach the optimal range of unit yield and selectivity.

2) A naphtha splitter, which was operated by the crude unit operations group, was operated at a low reflux ratio to improve variable cost performance. However, the reduction in fractionation efficiency resulting from this low reflux ratio affected the quality of the heavy straight run (HSR) and light SR (LSR) cuts, which affected

TABLE 2. Projected world oil demand, 2000–2030 Million 2010 2015

2000

2005

2007

LPG

6.1

6.6

6.9

7.3

Naphtha

4.2

4.9

5.2

Gasoline

19.5

21.2

21.8

Growth, %/yr 2007–10 2010–30

2005

Share, % 2007 2010

2020

2025

2030

2005–07

7.8

8.3

84.

8.2

2.5

1.9

0.6

8

8

8

2030

5.2

6.0

6.9

7.8

8.6

2.8

–0.1

2.6

6

6

6

8

21.9

23.1

24.1

25.1

26.6

1.5

0.1

1.0

25

25

25

25

8

Jet/kero

6.5

6.6

6.6

6.8

7.2

7.7

8.1

8.4

0.1

0.6

1.1

8

8

8

8

Gas/diesel oil

20.7

23.3

24.1

25.1

27.6

29.7

30.9

31.2

1.6

1.4

1.1

28

28

29

29

Fuel oil

10.4

10.0

9.7

9.4

9.2

9.0

8.8

8.8

–1.4

–1.1

–0.3

12

11

11

8

Other products

9.4

11.0

11.5

11.9

12.9

13.7

14.4

15.0

2.1

1.2

1.1

13

13

14

14

TOTAL WORLD

76.7

83.7

85.9

87.6

93.9

99.5

103.6

106.8

1.3

0.7

1.0

100

100

100

100

48

I JUNE 2009 HYDROCARBON PROCESSING

www.customidee.com

mangiarotti. looking ahead

mangiarotti at a glance Fabrication experience since 1930 Fabricator of large, high pressure, high temp heat transfer equipment, and heavy wall, large diameter, & exotic material process equipment.

Three fully staffed, equipped, modernized, & automated facilities.

One of our facilities is located in Porto Nogaro within 600 meters of the docks. All design, fabrication & testing are performed in-house. Weld up to 300 mm thick. Heavy items 1500 tons plus. Very long items over 150 meters long. Experience with an array of Process Licensor Technologies.

Engineer & design to internationally recognized codes & standards.

Highly qualified with an array of construction materials and advanced fabrication techniques. Certificates & Licenses including ISO-9001 & ASME Stamps. Experience with a multitude of international EPCs & end users.

Processing Equipment • Pressure Vessels & Reactors • Tubular Reactors • Ammonia Converters • Columns & Towers • Coke Drums • Separators • Splitters • Ammonia Synthesis Converters • Strippers • Regenerators • Contactors • Jacketed J k d Reactors R & Vessels • Hoppers • Scrubbers • Slug Catchers • Agitators • Crystallizers

Heat Transfer Equipment • S&T Heat Exchangers • Rod Baffle Exchangers • Internal Bore Welded • Finned Tube • HP/LP • Condensers • Steam Condensers • Gas Coolers • Waste Heat Boiler Packages • Ammonia Cartridges • Vertical Exchangers

January 2008 we acquired controlling interest of “Ansaldo Camozzi Nuclear Energy & Special Components” located in Milan, now Mangiarotti Nuclear.

Select 92 at www.HydrocarbonProcessing.com/RS Headquarters and Workshop Zona Industriale Località Pannellia, 10 33039 - Sedegliano Udine - Italy

Offices and Workshop Viale Sarca 336 20126 Milano - Italy

Heavy Equipment Workshop Z.I. Aussa Corno Via Enrico Fermi, 30 33058 San Giorgio di Nogaro Udine - Italy

Tel.+39.0432.918811 Fax +39.0432.918098

info@mangiarotti.it www.mangiarotti.it

SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION

Rigorous nonlinear unit simulator models allow the user to capture the nonlinearity of most refinery process units. Tuned models can be used to perform step-out evaluations for various parameters, making them particularly useful for opportunity evaluations.

• What is the key limitation in the distillate pool?

As indicated, investigating options for maximizing diesel is a complex issue. Several options are available. Unit test runs can be completed, but tracking all the changes across the entire facility is difficult, as is ensuring that production needs and specifications Flowsheet simulator. Another optimization tool is a fullcan be met when the changes are made. Using individual unit simrefinery flowsheet simulator, which typiulators is another approach. Although cally links several unit simulator models the model will provide information on together. If the full-refinery simulator ■ In today’s dynamic market, how a single unit will respond, capturcan propagate crude assay properties ing all the affects through the facility is and accurately simulate the complex refiners need a profit difficult, making this option marginreactions, separations and heat trans- enhancement process focusing ally better than test runs for significant fer effects of each refinery unit, then changes. complex refinery interactions can be on achieving sustained LPs and rigorous nonlinear simulamodeled seamlessly and accurately. The tions are the more likely candidates. The full-refinery flowsheet and the LP are profitability in a holistic fashion LP can capture the effects through the complementary tools. The full-refinery entire facility. However, it will likely not flowsheet easily handles the second and for the short- and long-term. capture many of the operational, equipthird order system interactions. An LP ment performance and catalyst effects can perform a similar function, but the structure has to anticinecessary to fully understand and evaluate the opportunities. pate the effects. As with LPs, rigorous simulators require periodic Maximizing diesel production is a classic application for a rigorous maintenance and substantial prep work for the set up. nonlinear simulation.

Under these uncertain times, refiners need profit enhancement Case study: Diesel maximization. Margins for diesel proteams for evaluating short- and long-term options for profitable duction remain strong, even in the current weakening economy. operations and regulatory compliance in a timely manner. Such Table 2 shows the projected world oil product demand through a team should be made up of key refinery personnel and external 2030.2 Over the next several years, gasoline demand growth will resources with a full array of expertise and tools at their disposal to be flat, while distillates remains strong, particularly through 2010. carry out effective evaluations after benchmarking and baselining In addition to long-term projections, gasoline vs. diesel selectivity is completed. Usually, an effective platform, such as an LP and/or varies seasonally, as illustrated in Fig. 2. Interestingly, the market a full refinery flowsheet simulator, is a key component for evaluatis showing a greater preference for distillates over gasoline. While ing various options. Finally, all such profit enhancement efforts others have examined the implications of choosing gasoline or must be examined within the confines of existing and expected diesel, the choice to increase diesel production is examined here to regulations imposed by government agencies for refinery operademonstrate the tools and methodologies discussed previously.3 tions and petroleum products. HP First, there are many potential opportunities to consider when LITERATURE CITED maximizing diesel production such as: 1 Calverley, S., “Refocusing on refinery profit,” Driving Competitive Advantage, • Are diesel component cutpoints at the right values?

Q2 2004.

• Are tray internals operating properly in crude distillation 2 Ohmes, R. and S. Sayles, “Analyzing and addressing the clean fuels and unit (CDU)/vacuum distillation unit (VDU), coke, and FCC? expansion challenge,” NPRA Annual Meeting, March 2007.

3 KBC Market Services, “World Long Term Oil & Energy Outlook,” November • Is diesel recovery being maximized out of gasoils?

2008.

• Do the diesel hydrotreaters have the capacity to handle additional material?

BIBLIOGRAPHY Bodewes, H., “The economics of residue processing—the Asia-Pacific context,” Hydrocarbon • Are the diesel range streams recoverable into the diesel pool?

1.85 1.80

Average 2008

1.75 Volume ratio

1.70 1.65

Asia, Jan/Feb 2006.

Haugseth, P. and G. Chukman, “Process profits with simulation,” Hydrocarbon Engineering, February 2005, pp. 57–60.

Jacobs, J., R. Ohmes and S. Sayles, “Gasoline or diesel,” NPRA Annual Meeting, March 2008, AM-08-59.

Lee, R., E. Leunenberger and R. Powell, “Optimizing the cat feed hydrotreater/FCCU complex with detailed simulation tools,” World Refining, July/August 2001.

LiveSmart BC Webpage, “President Obama Addresses Governor’s Global Climate Summit,” http://www.livesmartbc.ca/government/global_summit.

Polanco, D., “Monitoring and reducing a refinery’s carbon footprint,” NPRA Annual Meeting, March 2008, AM-08-41.

Westphalen, D. and H. Shethna, “Refinery wide simulation,” Hydrocarbon Engineering, March 2004.

1.60 1.55

Bill Fairleigh is a senior consultant with KBC Advanced Technologies, Inc. Parsippany, New Jersey. His primary responsibilities include flowsheeting and opportunity implementation for refinery clients worldwide. Prior to joining KBC, he worked as a revamp engineer for Koch Refining Co. in Corpus Christi, Texas, and as a refinery process engineer for Shell Oil in Houston, Texas, and Los Angeles, California.

1.50 1.45 Jan Jan Jan Feb Feb Mar Mar Apr Apr May May Jun Jun Jul Jul Jul Aug Aug Sep Sep Oct Oct Nov Nov Dec Dec Dec

1.40

FIG. 2

50

Seasonal demand for gasoline and distillate ratio: 2002– 2007.

I JUNE 2009 HYDROCARBON PROCESSING

Joseph Jacobs is a senior staff consultant with KBC Advanced Technologies, Inc. Robert Ohmes is an operations manager with KBC Advanced Technologies, Inc.

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PROCESS AND PLANT OPTIMIZATION

SPECIALREPORT

Rethink your liquid-liquid separations A fresh look investigates general principles in designing process coalescers R. CUSACK, Koch-Glitsch, LP, Wichita, Kansas

E

fficient liquid-liquid separations are an integral part of many industrial processes. Water entrained in oil and other hydrocarbon products can impede quality specifications.

Entrained oil in process water streams puts additional demands on effluent treatment systems.

Under the present economic environment, processing production demands push plants and existing equipment beyond the original design capacity. Equipment that formerly made clean-phase separations can no longer do so. And for new plants, strong economic incentives push process designers to incorporate equipment with smaller footprints to minimize weight, plot space and capital costs.

Several examples highlight typical issues encounter with liquidliquid separation operations and equipment applied in the hydrocarbon processing industry (HPI). Some general principles on design for liquid-liquid (and some vapor–liquid-liquid) separators will also be discussed Plant operators have found that efficient liquid-liquid separations are critical to achieve optimum plant performance But, how does one achieve this goal? What are the critical process parameters to be analyzed? What plant data need to be collected? What equipment options are available? How do you choose between these options? To begin the process of answering these questions, we must ask the fundamental question:

HOW WAS THE LIQUID-LIQUID DISPERSION FORMED IN THE FIRST PLACE?

Understanding how a droplet dispersion formed provides us with an idea of the size droplets present in the dispersion AND their relative volume. This information is crucial when creating effective designs for liquid-liquid separation processes. The three primary mechanisms for droplet creation are mechanical energy input, phase condensation or cooling and chemical reaction.

Mechanical energy input. The type of droplet distribution

produced by mechanical means is very dependent on the intensity and uniformity of shear forces present at the point of formation.

For example, a rotating impeller has the characteristics of high shear at the tips of the impeller and low shear at the center of the impeller. Result: It creates a wide range of droplet sizes while mixing the fluids. By contrast, if the fluid mixing occurs in an inline static mixer, the shear characteristics are uniform and produce a very well-defined, narrow droplet distribution. However, if these same liquids unintentionally mix via a centrifugal pump, the shear characteristics are typically very high and very non-uniform. The result is a very wide droplet distribution with a high percentage of very small droplets (often referred to as “fines”) that are much more difficult to separate.

Phase condensation or phase cooling. During phase condensation, droplets form on heat exchanger tubes (such as an overhead condenser on a distillation column) and then flow into piping and onto a separator drum. The droplet size distribution is large and easy to separate. By contrast, during bulk-phase cooling (such as in a product storage tank), water comes out of a solution and creates a dispersion that consists of much smaller droplets similar to a “fog.” Separating these droplets can be extremely difficult.

Chemical reaction. Similar to the situation of bulk-phase

cooling, chemical reaction creates a dispersion that has a very small droplet size distribution that is difficult to separate.

Separation curves. Fig. 1 shows the range of droplet sizes

associated with each of the listed mechanisms. This chart (Fig. 1) only illustrates typical ranges; it does not tell us the relative volume of the different size drops present in the dispersion.

An inlet volumetric frequency distribution curve (Fig. 2) presents the second piece of information needed for effective design. It shows the volume fraction of droplets in a dispersion as a function of droplet diameter. Three values of interest that provide the characteristics of the curve include: Maximum droplet size, d max Sauter mean droplet size, d 32 Mass mean droplet size , d 50.

These three values determine the shape of the distribution curve, which defines the degree of separation efficiency required to meet a particular outlet specification. The shape and limits of the inlet distribution curve are influenced not only by the mechanism of droplet formation, but also by the physical properties of the liquids and the characteristics of the system piping.

Mechanical Static mixers Mechanical agitator Centrifugal pump Two phase flow Condensation Heat exchanger from vapor From saturated liquid Chemical reaction 0.1

FIG. 1

0.0

10 100 Droplet size, μ

1,000

Typical droplet size ranges for various mechanisms (microns).

HYDROCARBON PROCESSING JUNE 2009

I 53

PROCESS AND PLANT OPTIMIZATION

Separator inlet distribution

0.007

Volume fraction frequency distribution, volume fraction per μ

Volume fraction frequency distribution, volume fraction per μ

SPECIALREPORT

d32 0.006 d50 0.005 0.004 Inlet distribution 0.003 0.002 0.001 dmax

0.000 0

Inlet droplet volumetric frequency distribution curve.

FIG. 2

120 d100

Removal efficiency, %

100

80

60

40

0.04

0.03 Outlet distribution 0.02

0.01 dmax

0.00 0

FIG. 4

20 Droplet diameter, μ

40

Outlet droplet distribution curve.

SEPARATING LIQUID-LIQUID DISPERSIONS

20

0 0

100

200 Droplet diameter, μ

300

400

Coalescer separation efficiency curve.

For example, if a large volume percentage of dispersed phase is present in a long run of piping between the points of mixing and of separation, a significant amount of coalescence can take place within the piping itself. Thus, the separation is easier. In this instance, the distribution curve shifts to the right. In contrast, if gas is present with the liquid in the pipeline and creates high velocity, the liquids will continue to mix as they flow through the pipe. This causes a more difficult separation, which shifts the distribution curve to the left.

The next piece of information necessary to analyze is from the efficiency curve of the separator (Fig. 3). This graph shows the removal efficiency as a function of droplet size. Of particular interest, the cutpoint droplet size, d 100, represents the smallest droplet 54

d50

that is removed at 100% efficiency. All droplets above this value are completely removed, and droplets below this value are removed with varying degrees of efficiency based on droplet diameter.

Changing the design of the vessel and/or internals can modify the shape of the separation efficiency curve, i.e. , shift the curve to the left or to the right as needed, to match separation requirements.

When the design meets the required process efficiency, then the outlet droplet distribution curve (Fig. 4) is the final curve generated. This curve represents the volumetric distribution of droplets remaining in the outlet stream when the coalescing and settling processes are complete. Armed with the physical properties of the liquids, the amount and size distribution of dispersion present and characteristics of the system piping and components, the next step is to create a design to separate liquids.

Separator efficiency

FIG. 3

d32

0.05

400

200 Droplet diameter, μ

Coalescer outlet distribution

0.06

I JUNE 2009 HYDROCARBON PROCESSING

The primary methods for separating liquid-liquid dispersions are: • Gravity settling • Enhanced-gravity settling • Coalescing • Centrifugal force • Electrical charge.

Separation designs frequently use the first three methods due to design simplicity, a wide range of applicability, and robust design.

The last two (centrifugal force and electrical charge) are specialized techniques that apply to limited applications, and several of those applications are switching to one of the other methods for the reasons noted above. We will focus only on the first three— gravity settling, enhanced gravity settling and coalescing.

Gravity settling. The simplest of all liquid-liquid separators is the gravity settler (Fig. 5). The gravity settler works solely on the principle of Stokes Law, which predicts the rate of rise or fall of droplets of one fluid inside another in accordance with Eq. 1.

The two most important physical properties in the settling process are illustrated by Eq. 1:

PROCESS AND PLANT OPTIMIZATION Feed

Light phase out

Eddy currents cause turblence

Chamber produces laminar flow and minimizes eddy currents

Reduced settling height enhances separation

Inlet dispersion

Force vectors working on drops

h

Force vectors working on droplets Heavy phase out FIG. 5

SPECIALREPORT

FIG. 6

Droplets reach liquid surface and coalesce

Outlet dispersion Outlet heavy phase

Separation action with enhanced-gravity settling.

Typical arrangement of a gravity settler.

Target wires or fibers

Stokes’ Law

Vs = where: Vs g ␳H ␳L d ␮

g ( H L ) d 2 18μ

(1)

Droplets impact on target

Some droplets are too small to be captured

Inlet dispersion

= = = = = =

Settling velocity, cm/s Gravitational constant, 980 cm/s2 Density of heavy liquids, g/cm3 Density of light liquid, g/cm3 Droplet diameter, cm Continuous phase viscosity, poise = g/cm-s

The key parameters are: • Density difference between the phases • Viscosity of the continuous phase.

Increased density difference makes the separation easier; increased viscosity of the continuous phase makes it more difficult. On the surface, the equation looks simple, straightforward, and easy to apply. However, this simplicity is deceiving because its formulation is based on some fundamental assumptions that limit its applicability.

The Stokes’ Law settling equation assumes 1) the drops are truly spherical and 2) they are moving freely in stagnant liquid.

Either of these assumptions is rarely the case (except possibly for the situation where the fluids are in a stagnant storage tank and subject only to gravity forces).

In more typical cases, the fluids flow horizontally in a vessel at Reynolds numbers well into the turbulent regime. This turbulence can deform the droplets from a spherical shape to an irregular shape that increases drag forces and hinders the settling rate. The turbulence can also create eddy currents that carry droplets along with them in the wrong direction. As a result, when sizing equipment using Eq. 1, it is necessary to apply “correction factors” that will account for the turbulence effects.

Determining these correction factors is very difficult to do even with today’s powerful computational fluid dynamics software.

For all the listed reasons, enhanced-gravity settling becomes the recommended settler type.

Enhanced gravity settling. As shown in Fig. 6, enhanced-

gravity settling minimizes the turbulence effects by dividing the stream into a number of separate channels. Dividing the flow into separate channels provides four primary benefits: • Decreases the effective diameter, thereby greatly reducing the

Droplets coalesce into larger droplets FIG. 7

Force vectors working on droplets

Coalescing process using surface energy forces with media.

Reynolds number of the flowing fluid and producing a deep laminar flow environment that enhances the gravity settling rate.

• Isolates the fluid in separate channels, thereby putting limits on how far droplets can “wander” and reducing the negative impact of eddy currents.

• Decreases the distance a droplet needs to rise or fall before reaching an interface, thereby greatly lowering the settling time requirement.

• Provides multiple interfaces inside the equipment where droplets can coalesce, thereby greatly increasing the coalescence process.

These benefits promote enhanced gravity settling and it is the recommended design rather than an empty settler vessel working on gravity alone. However, despite these improvements, there are situations where the droplets are so small, the density difference is so low or the continuous phase viscosity is so high that gravity forces alone cannot make an effective separation. In these cases, designs often add a coalescing step.

Coalescing. When gravity forces alone do not produce an efficient separation, adding a coalescing step (Fig. 7) to the process can improve separation efficiency. Coalescing designs insert targets into the flow path of the fluids. The droplets impact the targets and collect on the surfaces of the targets through surface energy forces. Once captured, these individual droplets combine (i.e., coalesce) to form larger droplets that are much easier to settle downstream.

HYDROCARBON PROCESSING JUNE 2009

I 55

SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION

■ It has long been said that “oil and water

do not mix.” But, actually they do. And when water and oil do mix, it usually results in major problems. For example, consider a car engine. A little bit of oil or a little bit of water in the wrong place can result in tremendous damages to the engine.

The efficiency of any particular coalescing medium can be related to the dimensionless Stokes number (St) shown in Eq. 2.

Target collection efficiency increases as the value of the Stokes number increases, i.e., the higher the value of the Stokes number, the greater the target collection efficiency.

Stokes number

St = Where: St = ␳D = v = d␳ =

D v (d ) 2

18μ d t Stokes number (–) Dispersed phase density, g/cm3 Superficial velocity, cm/s Particle diameter, cm

(2)

␮ = Viscosity of continuous phase, poise - g/cm-s dt = Target diameter, cm.

Based on Eq. 2, separator designs can increase target efficiency by increasing the velocity of the droplets, decreasing the continuous phase viscosity and reducing the target diameter.

Increasing the velocity increases the momentum of the droplets and causes more of the smaller diameter drops to impact the targets rather than flowing around them in streamlines. However, higher velocity at the same time increases the drag forces on the coalesced droplets. These forces can prematurely pull the droplets off the coalescing media; thus inhibiting the settling downstream within the settling zone. Consequently, there is an upper limit to this velocity. The goal is to find a balance between a velocity that optimizes impact on the targets and does not interfere with the settling of droplets in the downstream settling zone.

Decreasing the continuous phase viscosity reduces the drag forces as the droplets either rise or fall inside the continuous phase.

The lower drag forces allow the droplets to rise or fall at higher settling velocities, thus reducing settling time. The principle way to reduce the continuous phase viscosity is to increase the operating temperature. But there is a limit, because increasing the temperature also increases the mutual solubility of the phases, which inhibits separation. Again, designers must find a balance between the competing elements.

Decreasing the target diameter reduces turbulence around the targets and allows more smaller droplets to be carried into the targets by the flowing stream. In addition to size, the material of construction of the targets also plays a role. Materials that are preferentially “wet” by the dispersed phase make the best targets. For example, to coalesce water out of oil, hydrophilic (i.e., water-loving) materials, such as stainless steel or fiberglass, are often used. To coalesce oil out of water, oleophilic (i.e., oilFeed inlet

Light phase outlet

Inlet dispersion

Gauge glass

IC

Slotted pipe distributor

Select 164 at www.HydrocarbonProcessing.com/RS 56

Turbulence isolation plate

Interface Interface Liquid-liquid Heavy phase control separation media outlet

FIG. 8

Typical arrangement for a two-phase horizontal settler.

FIG. 9

Plate-pack or corrugated-plate separation media , which are used for higher separation efficiency.

PROCESS AND PLANT OPTIMIZATION loving) materials, such as fluoropolymers or polypropylene, are often used. And frequently, coalescing elements are made from a combination of both hydrophilic and oleophilic materials. The combination of materials actually enhances the overall coalescing efficiency above that of using either material alone. The reason that efficiency improves is not fully understood.

Once the target efficiency is optimized, selecting the depth of the coalescing element, as well as the number of targets per unit depth (i.e., the target density), will improve the overall coalescer efficiency.

Because of the complex nature for the coalescing process, it is impossible to effectively design separators from first principles alone.

Instead, the design must be practical based upon information gathered from either pilot-plant tests or similar application experience.

TYPES OF LIQUID-LIQUID SEPARATORS

Five basic separator types are applied in the listed liquid-liquid coalescing and settling internals: • Two-phase horizontal settler • Two-phase horizontal coalescer/settler • Three-phase horizontal coalescer/settler • Three-phase vertical separator • Two-phase vertical coalescer.

Which design is chosen depends both on the nature of the feed, as well as on the degree of separation required.

Two-phase horizontal settler (liquid-liquid). A twophase horizontal settler (Fig. 8) typically has a large percentage (5%–10% or greater) of the dispersed phase present in its feed. The most common location is downstream of the mixing or washing steps in a process such as caustic or water washing. Because of the large quantity of dispersed phase in the feed mixture, droplets can easily find each other and coalesce. The droplet size distribution tends to be large, with a Sauter mean droplet size in the range of 500 μ–1,000 μ. Unless the density difference is very narrow between the phases and/or the continuous phase viscosity is relatively high (i.e., greater than 5 cP–10 cP), the separation is usually not too difficult.

In these drums, the settling media typically is either plate or corrugated plate separator media internals (Fig. 9). These internals work on the principle of enhanced gravity settling. Plate-pack Primary coalescing media Feed inlet Inlet dispersion

Slotted pipe distributor

Turbulence isolation plate

Secondary “polishing” coalescing/settling media (optional)

Primary settling media

SPECIALREPORT

Separating the impossible—oil and water. Entrained

oil is found in several processing streams in the HPI. Here are some major areas that must handle entrain liquids: Refinery fractionator overhead reflux drums frequently become a bottleneck when the towers they service are fitted with higher capacity internals. The higher throughput rates in the modified tower increase phase separation demands on the overhead drum. Poor liquid-liquid separation in the drum creates high water levels in the reflux stream that can lead to corrosion on the trays and increased energy consumption in the tower.

Caustic and amine treaters are critical components in the light-ends section of any refinery. If the plant increases capacity, the treaters often experience increased solvent losses from the top of the treating tower. Any hydrocarbon entrained with the solvent from the bottom of the treating tower can cause foaming in the downstream stripping tower and this results in production losses.

Alkylation units in refineries have several steps where sulfuric acid, caustic or water mix with hydrocarbon streams and then are separated. It is critical that these separations are as sharp and clean as practical to minimize acid and caustic consumption as well as to minimize the potential for fouling/ corrosion in downstream equipment.

Wastewater treatment facilities are an integral part of any HPI facility. Upsets in the main production plant can often burden the water effluent treatment facilities with increased hydrocarbon removal requirements before water discharge. internals are the best choice when there is concern about fouling or plugging of the media from solids or other materials in the feed.

If the feed stream is clean, then a corrugated-plate media is the better choice; it provides higher separation efficiency per unit of length at a lower cost.

For a new installation, a slotted-pipe distributor introduces the feed to the drum and directs the inlet mixture at low velocity toward the upstream head of the vessel. This prevents turbulence effects at the inlet from affecting the rest of the drum. Installing a turbulence isolation plate between the inlet and the first elements of media is a recommended practice. The plate stops any turbulence created at the inlet before it can be transferred downstream where it could affect separation efficiency by sending “waves” down the drum.

Light phase outlet

Coalesced liquid Interface Gauge glass

IC Interface control Heavy phase outlet

FIG. 10

Typical arrangement of a two-phase horizontal coalescer/ settler.

FIG. 10A Isometric view of two-phase horizontal coalescer/settler.

HYDROCARBON PROCESSING JUNE 2009

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Vapor outlet

Vapor outlet

Mixed phase inlet Vane inlet devise

PROCESS AND PLANT OPTIMIZATION

Drain line

Mist eliminator Interface Level control IC

Gauge glass Turbulence Coalescing Settling Interface isolation media media control plate (optional) Heavy Light liquid liquid outlet outlet FIG. 11

Mist eliminator

LC

Mixed phase Inlet

Vane inlet device Vent pipe

Gauge glass Level LC control

Typical arrangement of a three-phase horizontal coalescer/ settler.

Light phase outlet Interface IC control Gauge glass

Liquid-liquid separation media Heavy phase outlet FIG. 12

FIG. 11A Isometric view of a three-phase horizontal coalescer/ settler.

While passing through the media, the droplet dispersion grows and settles. By the time it exits the media, it has coalesced into rivulets or streams. After exiting the media, these streams quickly rise and/or fall to the primary interface in the vessel.

The settling section of the drum is downstream of the media. In this section, the final settling takes place and control of the primary interface inside the drum occurs. Depending on process requirements, changing the position of the interface can optimize the settling zone performance. Raising the drum interface increases the residence time of the heavy phase. Conversely, lowering the drum interface increases the residence time of the light phase. However, at the same time, these changes also impact the distance the droplets must rise or fall before reaching the interface and coalescing. The position of the interface for optimum performance must strike a balance between these two different effects. Determining the optimum position is found through trial-and-error experimentation.

Two-phase horizontal coalescer/settler (liq.–liq.).

The feed stream for a two-phase horizontal coalescer/settler (Fig. 10) usually contains a much smaller percentage (less than 5%) of the dispersed phase. This is typically the case for streams coming from an upstream primary separator, condensers or coolers in the process, or streams going to/coming from storage facilities. The 58

I JUNE 2009 HYDROCARBON PROCESSING

Three-phase vertical separator.

droplet sizes are smaller than those for the previous case, usually having a Sauter mean droplet size in the range of 100 μ–300 μ.

Because of the lower volume of droplets present, they are less likely to find each other by random motion and to coalesce.

Therefore, separator designs for this case require a combination of both coalescing and settling media. As the flow passes through the first stage of media, the wires and/or fibers act as targets where the droplets collect and coalesce. As discussed earlier, adjusting the wire and/or fiber diameter, density, depth and material of construction controls the efficiency of the selected media.

Once the flow leaves the first-stage coalescing media, it typically enters a settling zone media where the coalesced droplets settle and form rivulets, similar to the process described for the two-phase settlers. The settling sections normally use corrugatedplate separators type media. For applications requiring very stringent outlet stream specifications (usually in the ppm range), installing a final stage of high-efficiency “polishing” media (and sometimes additional settling media) can effectively remove the last traces of very fine droplets. If there is a second “polishing” zone, then the process repeats with a first-stage coalescing media and then a settling media zone. Because the droplets are much smaller, the second zone uses a higher efficiency media.

As shown in Figs. 10 and 10a, similar to the situation with the two-phase settler, a slotted pipe distributor introduces the feed, and a turbulence isolation plate sits just upstream of the first coalescing media. The drum often uses a “boot” (or “hat”) to collect the small volume of dispersed phase and to serve as the liquid-liquid interface controller.

Three-phase horizontal coalescer/settler. In the HPI,

the three-phase horizontal coalescer/settler (Figs. 11 and 11a)

PROCESS AND PLANT OPTIMIZATION

SPECIALREPORT

is a common separator type applied in the vides the ability to remove the last remnants overhead systems of distillation columns or of entrainment. Again, leaving a bottom segas an upstream separator associated with oil ment open reduces plugging concerns. and gas production. As the name implies, Downstream of the media in the setthis vessel separates three phases: gas, lighttling end of the drum, an overflow baffle phase liquid and heavy-phase liquid (usutypically sets the liquid level of the light ally water). phase in the body of the vessel. A sepaThe volume of the dispersed phase presrate level control instrument regulates the ent in the feed can widely vary. Droplet light-phase liquid level on the downstream size distribution can also have a wide range side of the baffle. If a large amount of disbecause the feed-line flow is three phase persed phase is in the feed, then control and is flowing at a higher velocity due to of the interface is in the main body of the the presence of the gas. High velocity causes vessel on the upstream side of the overflow turbulence in the inlet line that can crebaffle (Fig. 11). If the amount of dispersed ate very small droplets, particularly if the phase is small, then typically adding a vapor-liquid surface tension and the liquid“boot” to the vessel provides the means to liquid interfacial tension are low. Therefore, collect the coalesced material and control proper design of the feed inlet device, gasthe liquid-liquid interface level. liquid separation media and liquid-liquid separation media are all critical. For the Three-phase vertical separator. feed inlet, designs often use a vane inlet Most vertical separators usually handle device that gradually reduces the momenonly two phases—liquid and gas. However, tum of the incoming stream to minimize if the liquid being separated forms two turbulence and droplet formation. Such phases, incorporating a liquid-liquid sepadevices can also acts as the first-stage gasration section in the bottom of the vessel liquid separation. is a common practice. Common applicaWhen space or mechanical issues prevent tions for these separators include pressureinstalling an inlet device, the next alternaletdown drums in refinery hydroprocessing tive is to direct the mixed flow toward the units and the bottom of quench towers in upstream head of the vessel. Because the ethylene plants. Figs. 12 and 12a show a entrance of the feed stream creates a large typical arrangement for the three-phase degree of turbulence, installing single- or vertical separator. double-turbulence isolation plates between As was the case for the three-phase the inlet and separation sides of the drum horizontal settler, introducing the feed to should reduce the effects of inlet turbu- FIG. 12A Isometric view of three phase the vessel is extremely important because vertical separator. lence on the downstream gas-liquid and impingement at high velocity on flat surliquid-liquid media. On some high-pressure faces can create fine droplets that are diffiapplications, particularly those related to upstream oil and gas cult to separate. Also, high vapor velocity across the liquid surface processing (may contain some solids), designs sometimes use a can re-entrain liquid up from the bottom of the drum. cyclone-type separator as the inlet device.

Using a vane-inlet device will gradually reduce the inlet For the gas-liquid separation in the top of the vessel, a vane-type momentum and evenly distribute the gas phase across the vessel mist eliminator is usually the best solution. The vane mist eliminadiameter. In the gas-liquid portion of the vessel, a wire-mesh mist tor is installed vertically in a housing within the vapor segment of eliminator provides high separation efficiency. In fouling and/or the drum with a drain that extends down into the liquid for sealing high liquid load situations where a wire-mesh mist eliminator purposes. In revamp situations where it is not possible to install the housing, the vane itself extends down into the liquid for sealing. In clean services, installing a section of wire mesh preceding the vane Light phase outlet provides an “agglomerator” stage that grows the droplets for easier Falling coalesced heavy phase removal by the vane, thus maximizing separation efficiency. droplets Coalescing For the liquid-liquid separation in the bottom of the drum, Falling drops media Interface the first stage is typically some type of enhanced-gravity separaGauge Interface tion media such as the plate-pack and corrugated-plate separators GG IC control IC GG glass described earlier. If fouling is a concern, then the plate-pack media Gauge glass Rising coalesced Rising Interface is the best choice. If the system is clean, then corrugated-plate sepalight phase drops rators media is the standard for the reasons given earlier. droplets Solids in the feed can present a very high potential for fouling, Coalescing which is often the case with separators in upstream oil and gas promedia duction. In these cases, a segment on the bottom of the media is left Heavy phase outlet open so that the solids can either pass through with the heavy phase FIG. 13 Typical arrangement of a two-phase vertical coalescer. or accumulate and be removed via flushing/drain nozzles. If very high separation is required, adding a second “polishing” stage proHYDROCARBON PROCESSING JUNE 2009

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SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION

cannot be used, designs often use a vane-type mist eliminator.

Sometimes a combination of vane and mesh is an option when the exiting gas stream requires very small amounts of entrainment.

In the liquid-liquid separation section in the bottom of the drum, either plate-pack internals (when there is a concern about solids or fouling), or corrugated-plate separators type internals provide efficient separation of the two liquid phases.

A level-control instrument maintains the gas-liquid interface in the main portion of the vessel. An interface control instrument located on the downstream side of the separation media maintains the liquid-liquid interface in the bottom of the vessel. The heavyliquid phase is withdrawn from the bottom of the vessel under the interface control. The light-liquid phase leaves the drum via an overflow nozzle. A vent line allows any entrained gas to rise to the top of the vessel.

Two-phase vertical coalescer. The two-phase vertical

coalescer (Fig. 13) is typically found in either the top or bottom of liquid-liquid contacting towers such as liquefied petroleum gas (LPG) amine treaters. These coalescers remove the very small amounts of entrainment coming from the interface and return them to the interface to minimize either the loss of valuable solvent or to prevent problems such as corrosion or off-spec product.

Media design for two-phase vertical coalescers is different from the horizontal design due to the flow direction for the coalesced droplets relative to the main process flow. In all the horizontal coalescers and settlers, the coalesced droplets rise or fall at an approximate 90° angle to the main flow. In a vertical coalescer, the coalesced droplets must flow at a 180° angle from

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the main flow, i.e., they need to flow countercurrent to the main flow direction or essentially “swim upstream.” Under these conditions, the droplets must grow to a larger size before breaking away from the media so that they can rise (or fall) against the drag force imposed by the continuous phase flow.

The droplet sizes to be removed in these applications are very small, with Sauter mean droplet sizes less than 100 μ. These critical applications require specially designed media. These elements have several zones and utilize a combination of knitted wire and fiber to achieve high efficiency and high capacity. Fig. 13 shows several options, based on the continuous and dispersed phases, for locating these coalescers in the tower relative to the location of the tower primary interface.

Overview. Efficient liquid-liquid separations is an integral part

of many HPI processes. The complex nature of both the creation and separation of liquid-liquid mixtures makes equipment design nearly impossible from first principles alone. Therefore, when addressing these processing problems, experience with similar applications or laboratory or pilot-plant data is needed. HP

Roger Cusack holds BS and MS degrees in chemical engineering from Manhattan College, New York. After graduation, Mr. Cusack worked for Exxon Research and Engineering Co., in the area of process design for both refineries and petrochemical plants. He subsequently worked for Chem-Pro Corp., the Otto York Co. and Glitsch Technology Corp., where he specialized in the areas of modular plant design and liquid-liquid extraction. Mr. Cusack is co-author of the liquid-liquid extraction chapter of Perry’s Handbook, Seventh Edition. At Koch-Glitsch, Mr. Cusack specializes in the areas of gas-liquid and liquid-liquid separation technology.

Select 165 at www.HydrocarbonProcessing.com/RS

PROCESS AND PLANT OPTIMIZATION

SPECIALREPORT

IT/automation convergence revisited Keeping automation close-coupled to operation is key A. G. KERN, Tesoro Corp., Los Angeles, California

I

n August 2008, two leading process control journals ran cover stories regarding the shared challenges facing information technology (IT) and automation, and of their coming “convergence�.1,2 As lead automation engineer at a major oil refinery, this caught me off guard because convergence wasn’t even on my radar.

After reading the articles more closely, the ideas behind convergence became more clear, and it is indeed a topic of high concern, though in a different light from the August articles. This article argues that what goes on between IT and automation is largely beside the point, while keeping automation close-coupled to operation is key.

Mind the gap (or “I’m with Cliff!�). The core challenge to any automation success, whether deployed in the automation or the IT domains, is to understand the opportunity being targeted from an operations point of view. If the end result is not a clear improvement in some aspect of plant operation, then no success has been achieved. Thus, the key area of interest is not between IT and automation, but between automation and operation.

It’s a mistake to assume that automation difficulties originate in poor IT/automation teamwork, or that capturing automation opportunities depends on extraordinary teamwork between these two. These groups naturally work well together. Both tend to be technically strong and well-focused on computer skills. The root source of most automation shortfalls is that both groups tend to be technically weak and poorly focused regarding plant operations.

In projects, IT often expects automation to serve as operation representative, but realistically, automation groups today are rarely able to fulfill this role effectively. Perhaps this is why problems seem to arise between IT and automation, while actually most gaps originate between automation and operation.

The best strategy to deal with the situation is to shore up operations knowledge of automation personnel. Many companies today retain the goal of pulling automation engineers from the ranks of successful operation engineers, but several factors have widely defeated this strategy. In particular, automation technology has greatly expanded and specialized, making it all the more difficult for individuals to bridge both roles. Moreover, modern focus on software means that many of today’s “automation specialists� are actually software engineers, not chemical engineers, and most have never worked directly in an operations environment.

To grok with operations personnel—with their workman’s motif, process lingo, knack for cooperation and constantly shifting priorities—is bridge building to be proud of, but few automation or IT personnel ever loiter here for long. The few individuals that do successfully bridge this gap can indeed play make-or-break roles in a company’s automation program. Unfortunately, most plants today have to persevere without the luxury of such people.

Of the many industry experts cited in the August articles, only Cliff Pedersen, a 39-year veteran of process control and oil refining, mentions the operation gap, rather than the IT/automation gap, as the key limitation. Convergence could work for some, he says, if you have an IT head who is knowledgeable about the manufacturing process, but most IT heads are computer smart, not plant smart.

Updated IT/automation paradigm. Fig. 1 is the traditional IT/automation paradigm, with problems presumed to arise from the undefined area of overlap. Indeed, many things are left undefined in this model, which is perhaps why it has historically led to finger-pointing rather than cooperation.

Fig. 2, on the other hand, introduces a number of important distinctions and brings a number of benefits: • Identifies the important and appropriate lines of communication • Provides a basis for proper design, selection and deployment of decision-support and automation applications • Maximizes productivity of specialized manpower in each area • Preserves integrity of the IT and (especially) automation domains.

Decision support systems (DSSs) are applications (properly) deployed in the IT domain that utilize information passed from the automation domain. Examples include data historians, data reconciliation, performance dashboards and (the business side of) alarm management, among many others. The primary users are process engineers, foremen, planners and managers (usually not operators).

The time scale of DSS applications is daily or weekly, so although these applications draw on control system information, they are not real time, don’t do control and don’t really belong in the automation domain. These applications are best supported by IT personnel, similar to other business-level applications. For most purposes, the IT/DSS group can interface directly with operations for design, Firewall

Corporate IT domain

Automation

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Traditional IT/automation paradigm.

HYDROCARBON PROCESSING JUNE 2009

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PROCESS AND PLANT OPTIMIZATION

Italian design A masterpiece

Firewall Corporate IT domain

Decision support

Automation (or process control)

Plant operations

Lines of communication t "QQSPQSJBUF CPVOEBSJFT BOE MJOFT PG DPNNVOJDBUJPO t "QQMJDBUJPOT EFQMPZFE JO BQQSPQSJBUF EPNBJO t &GmDJFOU NBOQPXFS VUJMJ[BUJPO FIG. 2

Creativity is the art we apply to achieve superior design and developments in technology. For over 70 years we have designed and supplied cost-effective technology, process plants and equipment for the oil & gas industry around the world. With our expertise we provide tailor-made solutions from studies and revamps to skid-mounted units and complete turnkey plants. Our own technologies are complemented by alliances with renowned licensors such as BOC, BP Amoco, IUT, WorleyParsons and UOP to provide state-of-the-art answers to design issues.

Oil & gas production facilities: separation, filtration, NGL and LPG recovery, stabilisation Gas & liquids treatment: amines, physical solvents, molecular sieves, iron oxide, glycol, silica gel, Merox™, sour water stripping Sulphur recovery: Claus, ammonia Claus, oxygen-enriched Claus, tail gas clean-up, Thiopaq™, redox, sulphur degassing, sulphur forming, advanced process controls Flue gas treatment: De-SOx, De-NOx & De-Dioxin, ammonia production Gas manufacture: low pressure gasification Special process equipment

deployment and support of such applications. Overlap between DSS and automation is primarily for setup of automated data transfer.

Process data historians have often followed this model, but the blossoming DSS field overall has ignored it, resulting in applications that often spill recklessly across domains, leading to chronic problems such as compromised integrity of the respective IT and automation domains, corruption of console operators’ scope of responsibility and work processes and stretching of valuable specialist manpower into each others’ areas of expertise.

DSS and automation specialists both have operations’ interests in common, but DSS specialists are computer engineers who put their operations knowledge to work in the IT domain, while successful automation specialists are usually chemical engineers who put their operations knowledge to work in the process automation domain. Making one from the other does not alleviate the manpower shortage that exists in both areas (especially in automation).

Making a DSS engineer from a successful automation engineer, a common enough mistake, is a net loss. And making automation engineers from computer engineers also has not panned out—computer engineers rarely learn to temper software results with the realistic process perspective necessary to make process automation work.

Reconcile Fig. 2 with your business. Achieving greater DSS

and automation effectiveness remains a crucial challenge, but “converging� IT and automation is not really indicated. The salient gap is in knowledge, not organization, and lays with DSS and automation already on the same side, but operations on the other. Today’s DSS and automation specialists are usually software oriented rather than operations oriented, with an idealized idea of operations born of the software that frequently forms a bigger part of their training and experience than actual operations. On the operations side of the gap is the complex reality of the operations environment. Reconciling Fig. 2 with your business, rather than convergence, can help bridge this gap to more skillfully and effectively capture the expected benefits of your automation and DSS programs. HP 1 2

SIIRTEC NIGI Engineering Contractors Via Algardi, 2 – 20148, Milan (Italy) Tel: +39-0239223.1 – Fax: +39-0239223.010 – E-mail: sinimail@sini.it – Web: www.sini.it Contact: marketing@sini.it

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Updated model provides many benefits.

LITERATURE CITED Policastro, E. F., “Plays well with others,� Intech, August 2008.

Katzel, J., “Automation, IT find teamwork pays,� Control Engineering, August 2008.

Allan Kern has 30 years of international process control experience and is currently working as a lead control systems engineer at Tesoro Corporation’s refinery in Los Angeles, California, USA. Mr. Kern is a licensed professional engineer, an ISA Senior Member and a 1981 graduate of the University of Wyoming.

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PROCESS AND PLANT OPTIMIZATION

SPECIALREPORT

Maximize ethylene gain in acetylene removal units New-generation catalysts and proper operating strategies offer improved selectivity and cycle length M. A. URBANCIC, M. SUN and S. BLANKENSHIP, Süd-Chemie Inc., Louisville, Kentucky; D. B. COOPER, Süd-Chemie Inc., Houston, Texas

A

troublesome byproduct from ethylene production is acetylene. It is a severe poison to downstream polymerization catalysts. Acetylene cannot be separated from ethylene by conventional distillation. Most acetylene converter units are tail-end hydrogenation units. Catalyst deactivation and nonselective hydrogenation with resulting ethylene loss are the most common operational problems associated with tail-end operations. Although proper operating strategies can minimize these problems, catalyst selection is the most important factor affecting unit performance. A new generation of selective hydrogenation catalysts can achieve significant gains in overall acetylene converter performance.

Ethylene production and purification. Steam cracking hydrocarbons is, by far, the most predominant method to produce ethylene. In this process, the feedstock is mixed with steam and raised to temperatures of 750°C–950°C for less than one second.

Due to the high temperature, large molecules are broken down into smaller ones, and saturated hydrocarbons become unsaturated. Result: A complex mixture of hydrocarbons is cooled, condensed and separated by downstream distillation.

One of the more troublesome byproducts of ethylene production is acetylene. It is problematic because it is a severe poison to the downstream polymerization catalysts used to produce polyethylene.

But acetylene cannot be separated from ethylene by conventional distillation. Depending on the feedstock and cracking severity, Ethylene overview. The annual global ethylene demand acetylene yield can be as high as 2.5 tons for every 100 tons of exceeds 100 million metric tons. As the leading commodity petethylene produced. Yet, the maximum specification for acetylene rochemical produced, ethylene is referred to as “the backbone of in purified ethylene can be as low as 1 ppm. the petrochemical industry.” The largest The most common process to remove single use of ethylene is for polyethylene acetylene from ethylene is by the use of a (PE) production. selective hydrogenation catalyst. While a ■ Choice of catalyst is the A key process for ethylene purificavariety of different selective hydrogenation is acetylene removal. This is nor- most important factor tion process configurations is available, mally accomplished with a selective the one that is still most commonly used hydrogenation catalyst. The tail-end affecting tail-end operations. is the so-called “tail-end” process. In this selective hydrogenation process was first process, the acetylene converter containintroduced in the early 1960s. Today, ing the selective hydrogenation catalyst this process still accounts for over 60% of all acetylene removal is located at the overhead of the deethanizer, after all of the units worldwide. hydrogen has been removed (see Fig. 1)—thus the term tail-end Profitable operation of a tail-end acetylene removal unit can be or back-end configuration. challenging. The most common operational problems with tailSince all of the hydrogen from the cracked gas is removed end units are 1) nonselective hydrogenation with resulting ethylene upstream, some hydrogen must be added back at the tail-end unit loss and 2) catalyst deactivation. The proper choice of operating to hydrogenate the acetylene. But the hydrogen/acetylene ratio strategy can help minimize these problems; however, the most must be carefully controlled to keep acetylene conversion and important factor affecting tail-end unit operations is catalyst. ethylene selectivity in proper balance. Less selective catalysts may Continued catalyst improvements over the years have allowed also require that carbon monoxide (CO) be injected at this point. plant personnel to make a remarkable shift in operating target Since CO is highly toxic, this can be a major safety concern. from low ethylene loss to high ethylene gain.2,3 To achieve maximum ethylene gain, it is critical to use the right catalyst for the Nonselective operation. Selectivity is simply the measureright feedstock. Recent successful commercial operation of the ment of ethylene gain across the converter. It is expressed as the latest generation tail-end catalyst has demonstrated that a further portion of acetylene that is converted to ethylene: increase of 20% (abs) in average ethylene selectivity or a doubling moles ethylene produced Ethylene selectivity = of the catalyst cycle length can be achieved. moles acetylene converted HYDROCARBON PROCESSING JUNE 2009

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high-hydrogen or high-temperature environment for an extended period, a change in the Quench/ catalyst can occur. At that point, it is seldom fractionate Compression Methanator possible to return to normal conditions. The H2S/CO2 overall selectivity will be decreased for the HC feed Demethanizer H2 Dryers remaining cycle. If the catalyst is inherently Cracking too active for the application, ethylene losses Compression furnace H2O will remain high despite attempts to optimize LP CH4 Cold box operating conditions.

Heavy gas + fuel oil Occasionally, a reactor may operate with such low selectivity (from zero ethylene gain to even a net loss) that the acetylene in the C2 fractionator Debutanizer Depropanizer Deethanizer C 2= feed cannot be reduced to levels below target.

Under these operating conditions, the reaction Tail-end of acetylene to ethylene is no longer substanMAPD C4 hydro. acetylene converter tially preferred and may be mass transfer limconverter ited. All hydrogen is consumed in the undeC3= C4 recycle sirable ethylene-hydrogenation reaction, and acetylene is present in the ethylene product.

C2 recycle The symptoms of this nonselective operaC5 plus tion include little to no hydrogen exiting the bed, excessive reaction occurring in the top FIG. 1 Flow scheme for a steam cracker showing location of tail-end acetylene removal unit portion of the bed, and acetylene slip from in the downstream effluent separation section. the reactor exceeding 2 ppm. The latter preWhen catalyst operates with low selectivity, ethylene losses occur, vents operators from feeding ethylene into a pipeline and may and profit is reduced. Often selectivity can be improved by making require flaring the product. If the only disposition for off-spec prodsimple changes in operating conditions, such as reducing excess uct is flaring, one off-spec incident that contaminates downstream hydrogen feed to the reactor, reducing temperature or tempering the equipment with acetylene can result in a loss of several hundred reaction with CO. However, if the catalyst has been operating in a thousand dollars in profit, not including environmental fines.

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SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION

Optimizing operating conditions early in the cycle and making small changes in conditions over time are paramount to maintaining good selectivity over the operating cycle and avoiding acetylene slip due to highly negative selectivity. However, the most important contributing factor to both good selectivity and achieving full acetylene conversion is choosing the correct catalyst for each application.

Right catalyst for right application. Not all tail-end acety-

lene conversion catalysts are created equal; thus, choosing the correct catalyst for each application is vital to a successful operation.

Factors to consider when choosing a catalyst include plant feed slate, reactor feed contaminants, reactor feedrate and future project implementation.

If the cracker process is mainly ethane/propane and the feed stream is relatively clean, a low-activity catalyst can be recommended. If the cracker processes naphtha, liquid feeds with potential metals contamination or olefinic streams from other sources (e.g., a nearby refinery) that contain low levels of contaminants, a more moderate activity catalyst is recommended. High-activity catalysts are generally only recommended when the reactor is required to operate at very high space velocities (such as in isothermal reactors) or in a high poison environment (including high CO that may enter with the hydrogen stream).

The main benefit of a moderate activity catalyst is flexibility.

When the operating facility is running at full rates or processing feedgas with contaminants, a moderate activity catalyst will have enough activity to provide full acetylene cleanup at reasonable selectivity. If market conditions require a change in feed slate or a cutback on rates, CO can be injected as a temporary poison to temper the

catalyst activity and increase overall selectivity. Ideally, a new generation catalyst would be flexible enough to allow rate and feed slate changes while maintaining high selectivity without CO addition.

Catalyst deactivation. There are three key mechanisms by

which selective hydrogenation catalysts deactivate in a normal operation cycle: 1) deposition of poisons from the feedstream, 2) sintering of the active metal particles and 3) formation of carbonaceous deposits on the catalyst surface.

Catalyst deactivation by poisons from contaminants in the feedstock is not widespread for tail-end catalysts. However, if it is occurring, a commercial poison guard can be used to maintain catalyst activity. Sintering (agglomeration of the active metal particles on the catalyst surface) is not a problem during normal use of tail-end catalysts due to the low operating temperature.

However, sintering can take place during the high temperatures experienced during catalyst regeneration.4 Both industrial experience and literature data indicate that sintered catalysts not only suffer a permanent activity loss due to diminished available metal surface area, but they also become nonselective. A sintered catalyst is also more vulnerable to metal poisoning, which can further accelerate catalyst deactivation. The potential to deactivate the catalyst by sintering can be reduced if the cycle length is extended, since the number of regeneration cycles is decreased. This can be achieved by applying new generation catalysts.

Due to the low hydrogen partial pressure in the inlet stream to the tail-end acetylene converter, polymer formation on the catalyst surface by oligomerizing acetylene is a normal occurrence.5 This condition is usually the major cause for catalyst deac-

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tivation. Under non-optimum operating conditions, polymer formation can occur quite rapidly.

Oligomers of acetylene that form on the catalyst surface can be divided into three groups, depending on how they affect catalyst performance. The first group can desorb from the catalyst, pass through the catalyst bed and be collected downstream as “green oil.” Green oil reduces the overall gain in ethylene selectivity. But it does not affect catalyst performance since it does not stay on the catalyst surface. The second group includes polymers with low H/C ratio that are left on the catalyst. The buildup of such polymers on the catalyst surface covers and reduces available active metal sites and results in activity loss. Also, accompanying the activity loss is a decline in the ethylene selectivity. The decline in selectivity is caused by the third type of oligomers that accumulate on the support. It has been suggested that the oligomers on the support increase the hydrogen “spillover” that accelerates the hydrogenation of ethylene to ethane.5 Also these oligomers can increase the internal diffusion limitation of acetylene, and thus increase the hydrogenation rate of ethylene to ethane at active sites on the internal surface of the pores. However, such diffusion limitation should not be a problem for catalysts that are properly manufactured with an “egg shell” metal distribution with less than 200 μm of palladium penetration.

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onstream (TOS) that is required to maintain a target acetylene conversion for the new and commercially available tail-end catalysts in a bench-scale test unit. By the end of the test, the new catalyst only required a temperature increase of about 13°C, while Catalyst A and Catalyst B needed increases of 53°C and 29°C, respectively. With the new catalyst, the deactivation rate is dramatically slower, thus allowing a much longer operation cycle.

Commercial plant operation has confirmed that the cycle length of the new catalyst can be two times longer than that of Catalyst B. Fig. 3 shows the decline in ethylene selectivity with TOS over the new catalyst and two reference catalysts in the same bench-scale test. The ethylene selectivity decline rate for the new catalyst was only half of that observed for the reference catalysts.

Assuming operation to a similar cycle length, it is clear that an operator will gain significantly more ethylene by using new generation catalysts. Depending on the actual operation cycle and EOR conditions, the average ethylene selectivity gain can be as high as 20% (abs).

The slower deactivation rate and excellent ethylene selectivity retention for the new catalyst is due to the strong resistance to polymer formation on the catalyst surface. Fig. 4 illustrates the significantly reduced polymer formation and higher ethylene yield over the new catalyst relative to Catalyst B. The lower polymer make also means lower greenhouse gas emissions upon regeneration.

lysts have operation cycles that are shorter than one year. For example, a start-of-run (SOR) selectivity of about 70% with rapid New tail-end catalyst: commercial decline to zero within four to six months experience. The first few commercial at end-of-run (EOR) is not unusual. To 60 significantly increase operation cycle length Catalyst A 50 and/or improve selectivity retention, a new 40 generation catalyst is highly desirable.

30 Catalyst B To achieve this goal, a new catalyst was 20 specially designed to reduce the polymer New catalyst 10 formation on the catalyst surface and 0 was tested extensively Time onstream in the laboratory.

Fig. 2 shows the FIG. 2 Increase in bed temperature over time onstream to maintain a target conversion for the new catalyst and two increase in bed temcommercial reference catalysts. perature over time Bed temperature - SOR temperature, °C

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SPECIALREPORT

New catalyst 20% Catalyst A Catalyst B

Relative selectivity, %

Ehtylene selectivity, %

First cycle on new catalyst

65% (abs) selectivity gain at 200 DOL

Multiple cycles on previous commercial catalyst 0

Time onstream FIG. 3

Decrease in selectivity over time onstream at constant conversion for new and two commercial reference catalysts.

Relative selectivity (new catalyst/catalyst B)

2.0

1.5

0.5

4 5

350

400

Commercial performance of new catalyst compared to previous catalyst demonstrates dramatically higher selectivity and cycle length.

Denise Branagan Cooper is currently a regional sales manager for petGreen oil

Ethane

Ethylene

The amounts of polymer, green oil, ethane and ethylene produced by the new catalyst relative to reference Catalyst B.

installations of any new catalyst systems are typically monitored extremely closely to confirm that the expected performance is achieved. Fig. 5 illustrates commercial data from one of the first users of the new-generation catalyst. The data shows that selectivity retention on the new catalyst far exceeds predictions and that cycle lengths increased by a factor of two under the newgeneration catalyst. More important, the operations levels were maintained without CO additions. With regard to regenerability, the new catalyst has already been regenerated in another commercial plant and was put back online with no apparent loss in activity or selectivity. HP

3

300

Süd-Chemie Inc. in Louisville, Kentucky. He has over 15 years of experience in catalyst preparation, testing and development. Dr. Sun earned a BS degree in chemistry from Nankai University in China and a PhD in technical chemistry from ETH Zurich Switzerland. Dr. Sun has done post-doctoral research in catalysis and reaction engineering at the University of Alberta and the University of Waterloo in Canada.

1.0

Polymer on catalyst

2

150 200 250 Days on load, (DOL)

Dr. Mingyong Sun is the R&D group leader for olefins purification catalysts with

0.0

1

FIG. 5

100

Dr. Michael Urbancic is the R&D Manager for petrochemical catalysts with Süd-Chemie Inc. in Louisville, Kentucky. He has over 30 years of experience in catalyst preparation, testing and development. During his 19-year career at Süd-Chemie, he has held a variety of positions in both R&D and technical services. Dr. Urbancic earned a BS degree in chemistry (with honors) from Purdue University and a PhD in inorganic chemistry from the University of Illinois.

2.5

FIG. 4

50

LITERATURE CITED Oil and Gas Journal, Vol. 106, No. 28, July 28, 2008, p 46.

Voight, R. W., J. S. Merriam and S. A. Blankenship, AIChE 8th Ethylene Producers Conference Proceedings, pp. 156–169, 1996.

US Patent No. 6,936,568.

Hall, J. B., B. J. Huggins, M. P. Kaminsky and B. L. Meyers, AIChE 6th Ethylene Producers Conference Proceedings, pp. 615–634, 1994.

Borodzinski, A. and G. C. Bond, Catalysis Reviews, Vol. 48, pp. 91–144, 2006.

rochemical catalysts with Süd-Chemie Inc. in Houston, Texas. She also serves as Süd-Chemie’s global topic expert in tail-end acetylene hydrogenation. Ms. Cooper worked as a process engineer for a global ethylene producer prior to joining SüdChemie three years ago. She holds a BS degree in chemical engineering from Virginia Polytechnic Institute and State University.

Steve Blankenship is employed with Süd-Chemie Inc. as a senior research scientist in Louisville, Kentucky. He has over 28 years of experience in the catalyst industry. Mr. Blankenship earned a BS degree in chemistry from the University of Kentucky in 1980.

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PROCESS AND PLANT OPTIMIZATION

SPECIALREPORT

Unifying framework for six sigma and process control The advances presented will improve quality and productivity P. B. DESHPANDE and R. Z. TANTALEAN, Six Sigma and Advanced Controls, Inc., Louisville, Kentucky

T

his article presents advances in six sigma and process control and illustrates how six-sigma concepts can be embedded in process control applications and vice versa. Thirtyfive percent of US businesses are reported to have embraced six sigma.1 The concepts presented could motivate many of the rest to embrace it as well. Doing so will further boost quality, productivity and competitive position. The advances reported may be used within process industries where the major impact factors are seen to be influencing both the mean and the standard deviation of response variables. Potentially, such applications exist across a wide array of diverse industries including polymerization, biological and biochemical processes, among others.

Background. The work presented was inspired by a handwrit-

ten page from the book, Jack: Straight from the Gut (Fig. 1).2 Taking “Order to Delivery Time” as an illustration, Dr. Welch made a convincing case for how focusing on improving the mean

performance (performance on average) is insufficient to compete in today’s global marketplace. One must focus on reducing variability, that is, the mean must be moved in a favorable direction and the standard deviation reduced. When this is achieved, all the benefits of six sigma accrue. To learn how, consider the relationship between the response variable, x, deemed to be normaly distributed, and the standard normal variable, z, for an application with a double-sided specification: USL μ

LSL μ z2 =

z1 =

(1a) (1b)

Here, μ is the response variable mean, ␴ its standard deviation, and USL and LSL are the upper and lower specification limits, respectively. The total area under the normal curve being one, the area between z1 and z2 represents good product while the area outside these limits represents defective product.

To reduce the defective area, the values of z1 and z2 must be increased. The specifications USL and LSL are driven by customer needs and are not amenable to adjustments for other reasons. To increase the “good” area, therefore, the denominator of Eq. 1 must be reduced and the numerator increased. That is, the standard deviation of the response variable must be reduced and its mean moved in a favorable direction (increased, reduced or brought to target, depending on the specific application under scrutiny). Fig. 2 shows the dramatic influence of increasing z on defect reduction.

160,000

Defects, ppm

120,000 80,000 40,000 0 1

FIG. 1

Mean performance versus reducing variability with permission of Jack and Suzy Welch from Jack: Straight from the Gut.2

FIG. 2

2

3

Z

4

5

6

Defects versus z (single-sided specification). For z = 4.5, defects are 3.45 per million opportunities.

HYDROCARBON PROCESSING JUNE 2009

I 73

PROCESS AND PLANT OPTIMIZATION

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I (first author) have stared at the handwritten page for several years ever since I acquired the book. His example implied that major impact factors exist that influence both the mean and the standard deviation of response variables. Indeed, GE six-sigma professionals showed that these factors do exist for both moving the mean to target and reducing the standard deviation to +1 day.

The example Dr. Welch used to make his point was clearly a transactional problem. I was telling myself, if major impact factors existed in transactional applications that influenced the standard deviations of response variables, not just their means, they must also exist in manufacturing applications. However, I could not understand how.

That is, how could the repetitive application of the same values of known major impact factors say, pH, temperature and time in an illustrative reacting system, give rise to varying values of the response variable (process outcome), say yield? We understand that the values of the response variable would not be identical time after time because of common-cause variability. Thus, if a number of experiments at the same factor levels were conducted, the response variables would not be identical due to common-cause variability arising from uncontrollable factors, but that is different from saying that the major impact factors influence the standard deviation response variables, not just their means.

To continue with the explanation, flows are routinely used in the process industries as major impact factors. If air-top pressure on a control valve was chosen as a major impact factor, then the associated response variable would or could be different for the same value of the air-top pressure because of line supply pressure changes. However, we also know how to overcome this lacuna by installing a fast control loop to maintain flow more or less constant regardless of changes in the supply pressure. Now, if flow were chosen as the major impact factor, we should get the similar values of the response variable for identical values of flow subject, of course, to commoncause variability considerations. Therefore, I was back to square one. Recently, the answer suddenly struck me. I have gathered sufficient courage to state it as a natural law: If identical values of major impact factors lead to substantially different values of the response variable(s) over and beyond common-cause variability considerations, then the population of response variable(s) must be heterogeneous or deemed to be heterogeneous.

Heterogeneity of response variable populations can occur for two reasons: (1) There are uncertainties in the inputs (impact factors)

that manifest themselves as heterogeneity of the response variable population and (2) the population of response variables is heterogeneous due to imperfect mixing in manufacturing applications. In either case, if a stratified random sample of size n representative of a heterogeneous response variable population is drawn and the mean and the standard deviation of the sample are computed, the major impact factors will be seen to be influencing both the mean and the standard deviation of the response variable. No mathematical proof of this law can be given, at least I do not know how. However, the way to disprove a natural law is to find evidence to the contrary. I look forward to reading about reader reactions to this natural law and to learning about any evidence that refutes it.

We cite two examples for further clarification of issues (1) and (2), one in nonmanufacturing and the other in manufacturing.

The nonmanufacturing example is intended to shed light on the uncertainty issue (1).

My associate, Mark Goldstein, a Certified Six Sigma Master Black Belt, uses a catapult experiment to demonstrate how the settings of major impact factors may be optimized to reduce variability. Here, several teams conduct experiments to show that identical values of major impact factors (e.g., pin location, launch angle and hook position) result in different values of the response variable (distance in feet the projectile travels). He also shows that the results of different teams are different owing to common-cause variability considerations. Mark then uses the data collected from designed experiments to develop two regression equations, one relating the response variable mean and the other, its standard deviation to the major impact factors. He then uses the regression models in an optimization algorithm to compute the best settings to apply for maximizing the mean and minimizing the standard deviation of the response variable.

He concludes the experiment by applying the optimized settings to the catapult and demonstrates that the variability has been reduced. In this application, though it is not obvious that the population of response variables is heterogeneous, we can only say that the response variable is deemed to have come from a heterogeneous population.

Here, the major impact factors (pin location, launch angle and hook position) appear to be deterministic. However, certain causes may manifest themselves as uncertainties in these major impact factors rendering them stochastic. Some examples of such uncertainties are: The pin over which the rubber band passes can rotate. The viscoelastic properties

PROCESS AND PLANT OPTIMIZATION of the rubber band can change over time. The duration for which the experimenter holds the launch position before releasing the projectile is not the same from one launch to the next, and so on. Because of these and other such uncertainties, the projectile may travel varying distances for the same values of major impact factors from one launch to the next.

Large batch or semibatch polymerization reactors are a good example of heterogeneity due to mixing issues in the manufacturing sector. Here, one or two samples are drawn from the sample port in the reactor vessel at the end of the batch cycle to infer product quality and the information is used in many cases to make changes in the major impact factors for the following batch in the belief that this will improve quality. Unfortunately, the population of response variable being heterogeneous because of mixing issues, the response variables may vary substantially from one location to another in the reactor. Therefore, had samples been taken from a different location in the reactor vessel, they would have suggested widely different changes in the major impact factors for the following batch.

To ascertain whether the variability in a response variable is due to common causes or uncertainties in the major impact factors or mixing issues, the suggested approach is to examine the regression equation relating the standard deviation of the response variable to the major impact factors. Highcorrelation coefficients (R 2 ) and favorable p-values are suggestive of uncertainties/ missing issues while poor-correlation coefficients and unfavorable p-values are suggestive of common-cause variability. Such situations are bound to exist in batch and continuous process applications as well.

What is the big deal with this natural law anyway? Well, in the past 35 years as a chemical engineering educator and industry consultant, my students and I developed a number of control laws aimed at achieving perfection. However, we never concerned ourselves (neither did anyone else to my knowledge) with the possibility that perhaps both the mean and the standard deviation of the response variable(s) can be improved.

Based on foregoing ideas, we believe we have advanced the state-of-the-art of six sigma and advanced control of static (batch) and continuous processes that should serve as a unifying framework for six sigma and process control.

Advances in process control of static systems. This advance pertains to

static systems such as batch polymerization

reactors. Leffew3 successfully tested constrained model-predictive control (CMPC) on a semibatch polymerization reactor for improved control of within the batch operations but the focus here is on batch-to-batch operations. For clarity, let us assume that a standard recipe-driven operating strategy is used. Thus, the focus in the proposed strategy is on major impact factors (e.g., initiator, modifier, chain transfer agent concentrations, etc.) to use from one batch to the next to achieve the best value of the mean of the response variable and the least possible standard deviation. The quality attribute of the product at the end of the batch cycle is the response variable. In theory, there could be more than one response variable.

The suggested approach is to conduct classical designed experiments in which the major impact factors are varied and the product quality ascertained at the end of each batch. The magnitude of changes in the factors must be large enough to produce sufficiently large changes in the response variable.

Since the population of response variables is heterogeneous, a stratified sampling plan must be used. The number of repetitions and replicates needed will depend on the heterogeneity of the population. This strategy will require a mechanism to draw random samples from various locations in the vessel and, therefore, could pose practical difficulties in some applications and will obviously incur additional expense. The financial benefits from improved product consistency, productivity and competitive position would have to justify the additional cost of multiple measurements involving repetitions and replicates. There is incentive here for companies to figure out how to make such measurements with instrumentation.

Once the experiments are devised and conducted, they will produce data that, when analyzed with standard statistical design of experiments software packages, will lead to two predictive equations per response variable: one for the mean and the other for the standard deviation, both as a function of the factors to be optimized. The terms involving products of major impact factors need to be retained in the regression model to account for the possible presence of interaction among them, something that chemical engineers typically do not do because it renders the model nonlinear. Quadratic terms could be included if warranted.

With the regression model at hand, one may proceed to compute the optimal values of the factors to apply to the process to achieve the best value of the mean of the response variable and lowest value of the Select 172 at www.HydrocarbonProcessing.com/RS 䉴

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PROCESS AND PLANT OPTIMIZATION

standard deviation. For an illustrative system with two factors and one response variable, the procedure is as follows: The regression equations selected for the illustrative example are of the form: y = a0 + a1 x1 + a2 x 2 + a3 x1x 2 (2a)

s = b0 + b1 x1 + b2 x 2 + b3 x1 x 2 (2b) The major impact factors, x1 and x2 , are in coded form. The terms are defined under Nomenclature at the end of the article.

The goal of optimization is to find the values of x1 and x2 such that a user-defined objective function is minimized. The complete optimization problem takes on the form: Min J = [mega{C1 (V1U +V1L ) +C 2 (V2U +V2L )} +C 3 x1 + C 4 x 2 ]

(3)

subject to the following constraints: a1x1 + a2 x 2 + a3 x1 x 2 y = a0 d 1

(4a)

b1 x1 + b2 x 2 + b3 x1 x 2 s = b0 d 2

(4b)

y + S1u V1 u = y u

(4c)

y S1L +V1L = y L s + S2u V2u = s u s S1L +V1L = s L

(4d)

x1 1

(4g)

x1 1

(4h)

(4e) (4f )

Select 173 at www.HydrocarbonProcessing.com/RS 76

Targets

+ –

Optimizer Eqs. 3 and 4

x1, x2

Compute ~– –~ y, s with Eqs. 2a and 2b

FIG. 3

– y, s

Process

–

+

CMPC system for the 2x2 example.

x 2 1

(4i)

x2 1

(4j)

An examination of Eqs. 4a–4j reveals that the terms on the right sides involve upper and lower bounds on y and s, the biases a 0 and b0 and the terms d 1 and d 2.

The bounds are user specifications and because the process regression model is known, the biases can be calculated from Eqs. 2a and 2b. The terms d 1 and d 2 are feedback signals, calculated by sampling the response variable at the end of each batch, computing its average and standard deviation, and subtracting from them the values ~y and ~s predicted from the respective regression equation. Chemical engineers will recognize this strategy as CMPC commonly used in continuous process applications. A block diagram of the CMPC strategy is shown in Fig. 3.

The linear objective function suggested in Eq. 3 is believed to be sufficient. The goal of optimization is to compute the best possible values of y and the least possible value of s such that the optimization index, J, is minimized. Due to the large size of mega (for example, 106 ), the optimizer will first attempt to eliminate the violation variables in Eq. 3.

If it can do so, only then will it focus on minimizing the costs as specified by the user. In other words, the bounds on y and s are treated as soft constraints while the bounds on x1 and x2 are treated as hard constraints and are therefore never violated. The terms C1 and C2 allow for relative weighting of the response variables while C3 and C4 allow for relative cost minimization of factors. Maximization may be achieved by making the signs of the cost coefficients negative. The user may specify the bounds on y and s or may specify targets. For the latter, the upper and lower bounds are set equal to the target. The number of response variables and the number of major impact factors will determine whether cost minimization (or profit maximization) is theoretically possible. This optimization problem is not amenable to linear programming due to the presence of the

PROCESS AND PLANT OPTIMIZATION product x 1x2 . This means that it may not always be possible to guarantee the global optimum. A local optimal solution should be sufficient.

Statistical packages have their own algorithms for solving optimization problems such as this. Here we describe a generic procedure for solving this constrained optimization problem.

1. Assume trial values of x1 and x2 within allowable ranges.

2. Compute y and s from Eqs. 4a and 4b, respectively.

3. Compute the violation and slack variables from Eqs. 4c–4f.

4. Compute J. If J has reached the minimum, end. If not, return to step 1 with new trial values of the factors and repeat.

Successful optimization strategies for this problem will ensure that the trial values of the major impact factors progressively move toward the optimum from one iteration to the next. The optimization surface in this problem and in problems like this is unusual owing to the presence of the product of the factors x1 and x2. This poses difficulties to the optimization algorithm in its ability to find the global optimum. A local optimum solution is believed to be sufficient.

Advances in six sigma. The foregoing ideas and concepts are an advance in the state-of-the-art of six sigma as well. The advance belongs to the “Improve Phase” of six-sigma investigations. In this phase, six-sigma practitioners traditionally utilize the regression models developed in the “Analyze Phase” to find the optimal values of major impact factors to obtain the best values of the response variable averages and the least values of their respective standard deviations. When this is done, defects reduce and the benefits of six sigma accrue. In the “Control Phase”, the response variables means and standard deviations are monitored typically with control charts to detect the arrival of new assignable causes. These assignable causes are worked on to return the response variables to their respective states of natural variability.

The foregoing concepts suggest that the Improve Phase should not be a one-time calculation in those applications where the major impact factors can be adjusted. When the major impact factors cannot be manipulated, the feedback signals d 1 and d 2 are zero. But when they can be manipulated, formulating it as a CMPC problem and using exactly the same procedure will lead to improved performance.

An experimenter conducts two3 full-factorial designed experiments on a process with three major impact factors and one dependent variable. The input–output data are analyzed with standard statistical software leading to the following regression equations: y = 1,051+ 163.8x1 + 301.6x 2 +75.07x 3 + 44.16x1 x 2

(5a)

s = 35.08 0.3951x1 +1.08x 2 + 2.21x3 +2.047x1 x3 + 2.685x 2 x3

(5b)

The problem is to determine the major impact factors x 1, x 2 and x 3 to apply to the process such that 1,025 ^ y ^ 1,050 and 32 ^ s ^ 34. Employing the optimization strategy described earlier, constrained optimization software gives the optimal values of the major impact factors according to: x1 = –0.751, x2 = 0.644 and x3 = –0.996. In the absence of modeling errors and load disturbances, these factors will result in y = 1,026 and ~s = 33 according to Eqs. 5a and 5b. The predicted values are within the specifications.

It is hoped that this article will motivate a number of companies to embrace the ideas presented and to find novel and higherimpact applications of the methodologies as suggested here. HP ACKNOWLEDGMENTS The authors thank Dr. Kenneth W. Leffew, DuPont Fellow, and business associate Mark Goldstein, Certified Master Black Belt, formerly of General Electric for their review and comments on this article.

a0 a1–a3 b0 b1–b3 C1, C2 C3, C4 d1 d2 J S s V xs

y

Advances in continuous process control applications. The concepts described here appear to have significant

potential in continuous process industries. The continuous process is assumed to operate under the command of a DCS system.

The approach to follow is to first determine the longest closedloop settling time of the process. This is the sampling interval for making changes to the major impact factors. Then, a host of major impact factors is identified; typically, they are the setpoints of feedback controllers or CMPC strategies. The response variables are quality attributes of the product. With this information at hand, classical designed experiments may be carried out involving repetitions and replicates to identify the major impact factors and regression models relating the response variable means and their respective standard deviations to the major impact factors.

Then, the foregoing optimization strategy may be implemented to achieve the best mean and lowest standard deviations of the response variables, the quality attributes.

Example. (Adapted from the Six Sigma Black Belt ParticipantGuide of Air Academy Associates, Colorado Springs, Colorado).

SPECIALREPORT

u L

1 2 3

4

5 6 7 8

NOMENCLATURE Bias in the regression model, Eq. 2a Regression coefficients in the model, Eq. 2a Bias in the s regression model, Eq. 2b Regression coefficients in the s model, Eq. 2b Weights on the response variable mean Cost coefficients on the major impact factors y (plant) – y (regression) s(plant) – s(regression) Optimization index Slack variables Standard deviation of the response variable, y Violation variables Major impact factors Response variable average SUBSCRIPTS AND SUPERSCRIPTS Upper limit Lower limit

LITERATURE CITED AND FURTHER READING Crocket, R. O. and McGregor, G., “Six Sigma Still Pays Off at Motorola, Business Week, December 4, 2006. p. 50.

Welch, J. and Byrne, J. A., Jack: Straight from the Gut, Warner Books, Inc., New York, 2003.

Leffew K. W., “Semi-batch Polymerization Process Control for Polymers Used in Microlithography,” presented at Center Jacques Cartier Conference on Modeling, Monitoring and Control of Polymer Properties, Lyon, France, 2007.

Burden, A. C., Tantalean, R. Z. and Deshpande, P. B., “Control and Optimize Nonlinear Processes,” Chemical Engineering Progress, 99, 2, February 2003. pp. 63–73.

Deming, W. E., “Quality, Productivity, and Competitive Position,” MIT Center for Advanced Engineering Study, Cambridge, Maine, 1982.

Deshpande, P. B., “Six Sigma for Karma Capitalism,” Six Sigma and Advanced Controls, Inc., Louisville, Kentucky, first quarter 2010.

Deshpande, P. B. Blog: www.on-a-quest-for-change.blogspot.com.

Deshpande, P. B., “A Small Step for Man: Zero to Infinity with Six Sigma,” HYDROCARBON PROCESSING JUNE 2009

I 77

SPECIALREPORT

PROCESS AND PLANT OPTIMIZATION

Six Sigma and Advanced Controls, Inc., Louisville, Kentucky, 2008.

Deshpande, P. B. and Tantalean, R. Z., “Process Control and Optimization,” Six Sigma and Advanced Controls, Inc., June 2007.

10 Deshpande, P. B., “Six Sigma Enlightenment,” Business World, October 4, 2004.

11 Deshpande, P. B., Makker, S. L. and Goldstein, M., “Boost Competitiveness with Six Sigma,” Chemical Engineering Progress, 95, 9, September 1999. pp.

65–70.

12 Deshpande, P. B., “Emerging technologies and six sigma,” Hydrocarbon Processing, April 1998.

13 Deshpande, P. B., “Globalization, Economic Development, and Competitiveness: Opportunities and Challenges,” R. N. Maddox Distinguished Lecture, University of Arkansas, Fayetteville, April 1998.

14 Deshpande, P. B., Ramasamy, S. and Yerrapragada, S. S., “Neural Nets Improve Batch Quality,” Control Engineering, April 1996. pp. 53–56.

15 Deshpande, P. B., Bhalodia, M. A., Caldwell, J. A., and Yerrapragada, S. S., “Should You Use Constrained Model Predictive Control?,” Chemical Engineering Progress, 91, 3, 1995. pp. 65-72.

16 Deshpande, P. B., Hannula R. E., Bhalodia, M. and Hansen, C. W., “Achieve Total Quality Control of Continuous Processes,” Chem. Eng. Progress, 89, 7, 1993.

17 Deshpande, P. B., “Improve Quality Control On-line with PID Controllers,” Chem. Eng. Progress, 88, 5, 1992.

18 Doyle, J. C., “Analysis of Controls Systems with Structured Uncertainty,” IEE Proceedings, Part D, No. 129, 1982. P. 242.

19 Harry, M. J. and Lawson, J. R., Six Sigma Productivity Analysis and Process Characterization, Addison-Wesley, Reading, Maine, 1992.

20 Krishnaswamy, P. R., Shukla, N. V., Deshpande, P. B. and M. N. Amrouni, “Reference System Decoupling for Multivariable Control,” Ind. Eng. Chem.

Research, 30, 4, 1991.

21 Ramasamy, S., Deshpande, P. B., Tambe, S. S. and Kulkarni, B. D., “Robust Nonlinear Control with Neural Networks,” Proceedings of the Royal Society, Series A, London, 449, June 1995. pp. 655–667.

22 Schmidt, S. R. and Launsby, R. G., Understanding Industrial Designed Experiments, Air Academy Press, Colorado Springs, Colorado, 1998.

9

Pradeep B. Deshpande is professor emeritus and a former chair of the Department of Chemical Engineering at the University of Louisville. He is also a visiting professor of management, Gatton College of Business & Economics, University of Kentucky, and the founder president and chief executive officer of the Louisville-based Six Sigma and Advanced Controls, Inc. (SAC). He was among the first to introduce six sigma training in corporate India and in engineering and MBA programs here, in Greece and in India. Dr. Deshpande is an author or co-author of six textbooks and over 100 refereed technical papers and presentations. During his 30 years on the faculty at the University of Louisville, he supervised 20 PhD graduates and over 40 master’s graduates. Dr. Deshpande is a fellow of ISA and a recipient of numerous awards for research and teaching. He is listed in Who’s Who in the World, and has taught at the Indian Institutes of Technology in Kanpur and Madras, University of Bombay Department of Chemical Technology and has spent a year on sabbatical at India’s National Chemical Laboratory in Pune.

Roberto Tantalean is a consultant with Six Sigma and Advanced Controls, Inc. He has served as assistant professor of chemical engineering at the Universidad de Trujillo for eight years.

Dr. Tantalean obtained his bachelor’s degree in chemical engineering from Universidad de Trujillo, a master’s degree in computer science from Universidad de Cantabria in Spain, a master’s degree and a PhD in chemical engineering, both from the University of Louisville, USA. He has developed and implemented a real-time communication facility for the constrained model-predictive controller software [ONLINE]t, and implemented an expert system for fault monitoring and abnormal situation management on an industrial-scale steam plant. Dr. Tantalean held an internship as a process engineer at Talara Refinery, Perú-Petro, the main Peruvian oil company. He has been part of the lecturer team in workshops in the areas of process control for Hindustan Petroleum Company in India and Six Sigma Training for the Private Universities Council in Kuwait. Dr. Tantalean is a consultant in the area of information technology for a US-based company He is a professor for a graduate school of engineering in a Peruvian private university in Software Engineering Quality.

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Designing and troubleshooting stabilization plant filtration systems—Part 1 Compatability of the filter cartridge media with the plant feed is a major consideration A. ATASH JAMEH, A. ZAMANI GHARAGHOOSH and S. BAZARGANI, Sarkhoon & Qeshm Gas Treating Company, Bandar Abbas, Iran; S. MOKHATAB, Gas Engineering Consultant, Vancouver, BC, Canada; and S. RAHIMI, Bid Boland II Gas Treatment Plant Project, Tehran, Iran

U

sually feed for a stabilization plant is supplied by a gas reservoir that has considerable dispersed brine water (formation water). Removing brine water from feed is essential in gas processing plants. One of the most common methods to remove water from unstabilized condensate is use of a filtration system such as a liquid/liquid coalescer. In selecting an appropriate filter cartridge element, it is important that its media be compatible with the plant feed. Installing a filtration system that includes a prefilter for the liquid/liquid coalescer could improve feed to stabilization plants. Filtration systems that have been installed in the Sarkhoon gas plant not only provide more stable plant operation, but also decrease corrosion in the debutanizer column and improve the quality of products such LPG and NGL. We have saved about $650,000 and reduced CO2 emissions by 3,466 metric tons annually.

In this part (1) we emphasize evaluating the filtration system and all the parameters that must be considered during design.

It is necessary to do a pilot field test to have a precise design for selecting filtration system cartridges and help the manufacturers develop their design criteria. The scopes of filtration system installation are: • Remove water and dissolved salt from fresh feed. Any change in temperature causes water to flash and salt deposits on the exchanger, control valve, and trays of the deethanizer and depropanizer towers.

• Remove solid particles and abrasive contaminants down to two microns from condensate to protect against downstream erosion and corrosion specifically of pumps and mechanical seals.

The package was installed in 2005. The prefiltration and coalescer systems use ultipleat high-flow and aquasep-plus cartridges, respectively (Fig. 1).

After installation and commissioning based on visual inspection (Fig. 2), plant laboratory results and operating experience, a large amount of water was removed from the prefilter and TABLE 1. Accumulated water in spherical storage when filtration system has been operated Water liters at LPG storage tank

Filtration remarks

02.03.05

20

Filtration in operation

12.03.05

4

Filtration in operation

12.04.05

2

Filtration in operation

19.04.05

50

Filtration in operation

08.06.06

32

Filtration in operation

86.4.18

40

Filtration in operation

14.08.07

450

Filtration out of service

17.08.07

350

Filtration out of service

19.08.07

300

Filtration out of service

22.08.07

250

Filtration out of service

Date

Condensate from upstream

S-701

Clear condensate to deethanizer F-701

FS-701

Background. The stabilization plant (Unit 700) of the Sark-

hoon refinery has been designed for stabilizing 1,550 tpd of unstabilized condensate. The products of this unit are 11,940 bpd of LNG and 1,048 bpd of LNG at standard conditions. To remove water and dissolved salt from the condensate feed to the stabilization plant, the liquid/liquid coalescer system was designed to continuously separate solids and water from the condensate feed to the stabilization plant down to two microns absolute of water.

Brine water FIG. 1

Brine water

Filtration system schematic (surge drum, prefilter and liquid/liquid filter coalescer).

HYDROCARBON PROCESSING JUNE 2009

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MAINTENANCE/RELIABILITY

FIG. 4

Agglomeration of black/brown-colored deposit on the upstream layer of the media.

they could mainly be metallic particle sands and some unknown nonmetallic particles like gel or polymer compounds.

FIG. 2 Samples from filtration system inlet (left) and outlet A 20-micron prefilter and phase separation cartridge were (right). installed. According to the coalescer performance results obtained from field test, it is believed that this change did not affect the coalescer condensate feed. At the filtration system outlet the global separation performance. samples were very clear and sparkling.

After four days the new prefilter cartridges were blocked, which There was a big difference in water concentration in LPG is not acceptable to the cartridge manufacturer due to increasing before and after coalescer installation due to the high-efficiency pore size from 2 to 20 microns. separation in the coalescing section (Table 1).

By receiving data from the plant it was clear that each cartridge An ultipleat high-flow filter element is used in F-701 to remove has about 1,200 grams extra weight. This indicated that the filter solids from the condensate to protect the aquasep liquid/liquid isn’t blocked due to solids and contamination size. coalescer elements. Prefilter lifetime was lower than expected (the Loading is not an issue for rapidly increasing pressure drop cartridges blocked every 10 days), because it has been shown these in normal operation. In amine filtration each ultipleat high-flow elements have three months lifetime in some references. cartridge can have up to a 12 Kg weight difference between clean and clogged conditions.

Root-cause survey and solution. To determine the cause We believed that after confirmation the amount of contaminants of the prefilter‘s cartridge blockage a used filter cartridge was sent that the prefilter had removed, as the weight data provided by the to a scientific laboratory for examining particle size distribution plant would indicate, the solids content would be quite low. and specifying the nature of the solids.

To specify which causes were the main issues, the following paramAfter performing a complete laboratory test, we submitted a eters had to be checked to be sure of the system’s designed values: complete report comprising detailed information about particle 1. Review the prefilter design regarding the number of elesize and nature. Microscopic observation of collected contamiments, clean pressure drop, filter media compatibility with the nation in the used cartridge revealed that the contamination is process stream and removal rating. mainly made of black-colored particles with an average size of 2. Check the prefilter pressure vessel mechanical design 10–40 μm. However, very fine particles smaller than 5 μm are hydraulics. present as well (Figs. 3 and 4).

3. Check the gas trap and condensate evaporation in the filter The exact nature of the contaminants was confirmed in the pressure vessel to high volatility. laboratory with an elemental analysis using a scanning electronic The scientific laboratory of the cartridge manufacturer microscope coupled with energy dispersive spectrometry. Based approved the prefilter design, number of elements and compaton observations by using the optical microscope, it is clear that ibility. The actual clean pressure drop is less than 25 mbar at design flowrate.

Although we couldn’t believe that the high-flow element was completely compatible with our plant, it is not discussed in this article. That subject will be discussed in part 2.

The scientific laboratory also approved the prefilter mechanical and internal designs such as the tube sheet and sealing mechanism based on approved drawings.

For a gas trap upset the condensate Rvp FIG. 3 Different layers of the upstream (left) and downstream sides (right) of the media is very low and due to the inlet/outlet nozzle pack structure. tube sheet and filter media pressure drop, 80

I JUNE 2009 HYDROCARBON PROCESSING

MAINTENANCE/RELIABILITY

Sarkhoon gas wells (sour gas) Corrosion inhibitor

Inlet separation (2-phase separators) (Unit 200)

Sarkhoon gas wells (sweet gas) Corrosion inhibitor

Sour gas

Sweet gas Condensate

Glycol regeneration (Unit 600) Corrosion inhibitor

FIG. 5

The gel layer is between the hydrocarbon condensate and brine water.

FIG. 6

Rich glycol

Lean glycol

Gas sweetening unit (Unit 1,000) Corrosion inhibitor Antifoam Ethanolamine (DEA)

Stabilization unit (Unit 700)

Dehydration unit deethylene glycol (Unit 400) Dry gas to sale

Chemical injection at the plant.

There are two main sources for gel formation: • Inside the refinery such as in preliminary separation, sweetening, dehydration and NGL recovery • Upstream processing where the corrosion inhibitor is injected by the exploration company in gas well production.

Regarding the first item, the major vessel that the probability of the gel source might have come from was selected. Samples taken from the equipment (Fig. 7) proved that the FIG. 7 Sample from surge drum S-701 (left). Gel produced in the laboratory on dosing 0 ppm, 50 ppm, 100 ppm, 200 ppm and 500 ppm (from left to right). process plant could not be a main cause of gel formation. Samples showed the amount some evaporation will occur inside the vessel and it seems that a of gel that could be collected at the inlet separators is more than gas trap forms inside the vessel. We have checked gas bubble traps had been taken from the dehydration unit’s separators. So this that are able to provide an artificial clogging of filter media by could have enhanced the hypothesis that the main cause of gel providing complete prefilter vessel ventilation. This case has not source formation should be the upstream wells. As a result of been considered completely by the cartridge manufacturer yet. finding the root cause of the problem all activities concentrated Based on all previous testing on filter media, liquid condensate on upstream well production and all the chemicals that were being fluid and the site test on the prefilter housing, we conclude and injected into the wells to protect against corrosion. believe that there are unknown components in the condensate stream For the second item, gel formation and its delivery to gas plants that generated from the process upstream. In the plant laboratory we should be controlled and reduced by the exploration company have carried out visual checking of the unknown components. after having a technical discussion with the corrosion inhibitor The sample that has been taken from the surge drum boot (Fig. manufacturer about compatibility of the corrosion inhibitor 5) is gray and loaded with black solid particles. with hydrocarbon condensate. We made a decision to replace After a short settling time, the sample separated into three the chemical with a new one. All of the operating parameters phases. The top phase looked like condensate oils while the botrequired to evaluate the effect of the new chemical on the cartom phase is brine water and the middle phase is milky/grey. tridges were controlled. We checked the effect of this chemical These gel-like components are likely made of combination of by taking samples from surge drum S-701 and the condensate some hydrocarbons with none of the partially soluble components quality at coalescer FS-701’s exit was acceptable. As a result the being corrosion inhibitor. final decisions were: Gel has a blocking effect on the filter media and various chemi• Remove gel produced by filtration upstream of the stabilizacals are injected into wells upstream for process and maintenance tion unit operation. When a sample is shaken, we clearly observe some vis• Optimize gas processing and control chemical qualities, rate cose gel-like component on the walls of the bottles. It seems that and dosage. the main chemical that is likely present in the condensate is an Prefilter F-701 was removing all impurities and solids from the organic-oil-soluble anticorrosion inhibitor that is injected into the hydrocarbons so the emphasis on this subject by the contractor gas wells. However, according to the drawings provided, the conis acceptable. densate stream collected from the two-phase separator is mixed with Removing gel by installing a new filter was presented by the filter a hydrocarbon stream collected from NGL recovery and the glycol designer, but after pilot testing it was clear that applying this policy dehydration unit upstream of the stabilization unit (Fig. 6). could be a final solution for filtration. Then we concluded that HYDROCARBON PROCESSING JUNE 2009

I 81

MAINTENANCE/RELIABILITY another option must be selected. Regarding the second option, we changed the corrosion inhibitor and controlled the injection. After a while it was determined that changing the chemical improved filtration system performance but other issues were raised that will be discucssed in the next article by the authors. HP ACKNOWLEDGMENT The authors appreciate the efforts of the following individuals and departments for their persistence in helping us to collect the plant operation data.

Thanks to M. Mihandost of production, M. Nori of laboratory and M. Salehadai of Panid Co.

1 2

3

LITERATURE CITED Hawn, R. R., Ellington, E. E., et al, “International Gas Processing Prospects Look Bright to 2000,” Oil & Gas Journal, July 20, 1992.

Pauley, C. R., Langston, D. G. and Betts, F. C., “Solving Foaming and Amine Loss Problems Treating Plant,” presented at the AICHE Summer National Meeting, San Diego, California, August 1990.

Murphy, W. L., “Practical In-Service Simulation Tests for Rating of High Aerosol Coalescing Performance,” PEDD-FSR-101a, Pall Corporation Equipment Development, November 1984.

Ahmad Zamani Gharaghoosh is the head of technical inspection department in Sarkhoon & Qeshm Gas treating company (SQGC). He joined National Iranian Gas Company ( NIGC) In 1997.

Mr. Zamani Gharaghoosh has more than 11 years of experience in gas refinery processes, static equipment inspection, risk-based inspection and is a corrosion specialist at the gas refinery.

Saifollah Bazargani is head of technical services and engineering in Sarkhoon and Qeshm Gas Treating Company, Iran. He joined the National Iranian Gas Company (NIGC) in 1996 and has extensive process engineering experience in separation processes and acid gas removal. Mr Bazargani has also worked with National Iranian Petrochemical Company for seven years in the olefins Plant. He holds a BS degree in petrochemical engineering from Amir Kabir (Ploytechnic Tehran) University of Technology in 1983.

Saeid Mokhatab is an internationally recognized expert in the field of natural gas engineering with a particular emphasis on gas transmission, LNG and processing. He has been involved as a technical consultant in several international gas-engineering projects and published over 150 academic and industry-oriented papers.

Mr. Mokhatab is a member of the editorial board for most professional oil and gas engineering journals, and serves on various SPE and ASME technical committees.

Abolfazl Atash Jameh is the head of process engineering, a division of engineering and technical services, in Sarkhoon and Qeshm Gas Treating Company, Iran. He joined the National Iranian Gas Company (NIGC) in 1999 and has 10 years of experience in process engineering, modeling and optimization as well as troubleshooting gas processing units. Mr. Atash Jameh holds an MS degree in chemical engineering from the Sharif University of Technology in 1998 and a BS degree in chemical engineering from the Petroleum University of Technology in 1995. He has authored and coauthored more than seven papers in national and international conferences.

Samad Rahimi is the head of engineering and design for the Bid Boland II Gas Treatment Plant basic design project. Before that, he was the head of engineering and technical services in Sarkhoon and Qeshm Gas Treating Company, Iran. His areas of specialty include gas processing, LNG separation and hazard identification.

Mr. Rahimi joined National Iranian Gas Company (NIGC) in 1989 and has held various design engineering positions including chief engineer for NIGC. He received his BS degree in chemical engineering from the Petroleum University of Technology.

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Minimize vaporization and displacement losses from storage containers Consider using this new calculation for recovery A. BAHADORI, Curtin University of Technology, Perth, Western Australia

A

storage tank is generally not pumped completely dry when emptied. The vapor above the remaining liquid will expand to fill the void space at the liquid’s vapor pressure at storage temperature. As the tank fills, vapors are compressed into a smaller void space until the set pressure on the vent/relief system is reached. Some filling losses are associated with the liquid expansion into the tank.

Vapors emitted from a storage tank’s vents and/or relief valves are generated in two ways: • Tank vapors forced out during filling operations (displacement losses) • Vapors generated by liquid vaporization stored in the tank.

Storage classification. There are two types of storage clas-

sifications, above ground and underground. Categories include: • Atmospheric. These tanks are designed and equipped for content storage at atmospheric pressure. This category usually employs tanks of vertical cylindrical configuration that range in size from small shop-welded tanks to large field-erected tanks.

Bolted tanks and, occasionally, rectangular welded tanks are also used for atmospheric storage service.1 • Low pressure [0 to 17 kPa (ga)]. These tanks are normally used in applications for storage of intermediates and products that require an internal gas pressure from close to atmospheric.

The shape is generally cylindrical with flat or dished bottoms and sloped or domed roofs. Low-pressure storage tanks are usually of welded design. However, bolted tanks are often used for operating pressures near atmospheric. Many refrigerated storage tanks operate at approximately 3.5 kPa (ga).1 • Medium pressure [17 to 100 kPa (ga)]. Medium-pressure tanks are normally used for the storage of higher volatility intermediates and products that cannot be stored in dished bottoms and sloped or domed roofs.1

• High pressure. These tanks are generally used for the storage of refined products or fractionated components at pressures above 100 kPa (ga).1 • Underground. Gas processing industry liquids may be stored in underground, conventionally mined or solution minedcaverns. No known standard procedures are available for this type of storage; however, there are many publications and books covering the subject in detail.1 Displacement losses. The combined loss from filling and emptying is considered a working loss or displacement loss.

Evaporation during filling operations is a result of an increase in the liquid level in the tank. As liquid level increases, the pressure inside the tank exceeds the relief pressure and vapors are expelled from the tank. Evaporative loss during emptying occurs when air, drawn into the tank during liquid removal, becomes saturated with organic vapor and expands, thus exceeding the vapor space capacity.

Vaporization losses. Vapors are generated by heat gained

through the shell, bottom and roof. The total heat input is the algebraic sum of the radiant, conductive and convective heat transfer. This type of loss is especially prevalent where light hydrocarbon liquids are stored in full pressure or refrigerated storage. This is less prevalent but still quite common in crude oil and finished product storage tanks. These vapors may be recovered by using the vapor recovery system. To calculate vaporization in tanks, the sum of radiant, conductive and convective heat inputs to the tank must be taken into account.

Approximate vapor losses in kg/h can be calculated by dividing the total heat input by the product latent heat of vaporization at fluid temperature.1

TABLE 1. Tuned coefficients used for polynomial models. x

Ax

Bx

Cx

1

1.96692469015064E-1

–2.997855656735E-3

1.5038580989E-5

2

–3.6068872437314E-2

3

1.725348224825E-3

4

–1.26233316566389E-5

5.96919209298E-4 –2.5923852687E-5 1.833674141261E-7

–3.065251613E-6 1.30357956E-7 –9.137656612E-10

Dx –2.4675901E-8 5.109185E-9 –2.174E-10 1.5231731E-12

New proposed correlation. Eq. 1 presents a new correlation where four coefficients (a, b, c and d) are used to relate the filling losses from storage containers in percent of liquid pumped in with working pressure in kPa(abs). However, vapor pressure at liquid temperature kPa(abs) must be considered when estimating the displacement losses from storage containers.

HYDROCARBON PROCESSING JUNE 2009

I 83

GAS PROCESSING DEVELOPMENTS loss = a = bP = cP2 = dP3

where a, b, c and d are derived through polynomial equations in the third order (A x = Bx Pv = C x Pv2 = Dx Pv3 ) with x denoting either 1, 2, 3 or 4. The tuned coefficients (A x, Bx, C x and Dx ) used in polynomial models are given in Table 1 which covers reported data for working pressure less than 250 kPa (abs) and for vapor pressure at liquid temperature less than 100 kPa (abs).

Results. Fig. 1 compares the new proposed correlation results

for predicting the filling losses from storage containers in percent of liquid pumped in, to the available reported data. As shown, there is an agreement between predicted and reported values. This correlation covers the reported data for working pressure less than 250 kPa (abs) and for vapor pressure at liquid temperature in kPa (abs) for less than 100 kPa (abs). HP A B C D Loss P Pv

1

NOMENCLATURE Coefficient Coefficient Coefficient Coefficient Filling loss in percent of liquid pumped in Working pressure in kPa (abs) Vapor pressure at liquid temperature in kPa (abs)

LITERATURE CITED Gas Processors and Suppliers Association Data Book, 12th Edition, Tulsa, Oklahoma, 2004.

Filing loss, percent of liquid pumped in

(1) 0.30 0.25 0.20 0.15

Working P=5 kPa (g) Data Working P=10 kPa (g) Data Working P=25 kPa (g) Data Working P=50 kPa (g) Data Working P=100 kPa (g) Data Working P=150 kPa (g) Data

0.10 0.05 0.00 10

FIG. 1

20 30 40 50 60 70 80 90 Vapor pressure at liquid temperature, kPa (abs)

Comparison of predicted and reported data of the displacement losses from storage containers.1

Alireza Bahadori is a PhD student in the chemical engineering department at the Curtin University of Technology, Perth, Western Australia. Previously, he worked as a senior process engineer at National Iranian South Oil Company (NISOC) for 10 years and was involved in several large-scale oil and gas projects. Mr. Bahadori is the author and co-author of 70 referred journal papers and the recipient of the Australian Government’s Department of Education Science and Training Endeavor International Post Graduate Research Scholarship (EIPRS). He’s also received the state of Western Australia’s top scholarship through the Western Australia Energy Research Alliance (WA:ERA). Mr. Bahadori is a member of Engineers Australia.

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Receive Monthly e-newsletters providing an early preview of upcoming special editorial features, which provide operational and technical insights: August 2009: Fluid Flow and Rotating Equipment • Pumps, valves and compressors • Level gauging • Custody transfer systems–flow meters, process analyzers • Tankfarm management September 2009: Refining Developments • Product/cracking methodologies, lowsulfur, low-aromatic transportation fuels • Hydrogen management • Catalyst developments • Environmental compliance • Alternative/Biofuels October 2009: Process Control and Information Systems • Analyzers–online/portable/benchtop • Process control systems and components • Fieldbus systems and networks • Instruments–transmitters, analyzers, controls, meters

2009

HPI Construction Boxscore Update

JUNE

COUNT OF NEW PROJECTS YEAR AGO (JUNE 2008) Rfg United States

P’chem

Gas

THIS ISSUE (JUNE 2009) Other

Total

Rfg

P’chem

Gas

Other

Total

40

13

15

5

73

United States

5

1

3

1

10

6

6

7

4

23

Canada

3

1

1

0

5

Latin America

10

10

5

3

28

Latin America

12

7

7

0

26

Europe

13

19

27

22

81

Europe

15

11

17

6

49

Asia/Pacific

33

30

29

10

102

Asia/Pacific

14

13

10

1

38

4

5

4

7

20

Africa

2

3

0

1

6

16

19

13

10

58

Middle East

3

9

3

2

17

122

102

100

61

385

54

45

41

11

151

Other

Total

184

104

714

Canada

Africa Middle East Total

Total

COUNT OF TOTAL ACTIVE PROJECTS YEAR AGO (JUNE 2008) United States Canada

THIS ISSUE (JUNE 2009)

Rfg P’chem

Gas

Other

Total

239

168

185

79

671

United States

Rfg

P’chem

254

172

Gas

82

22

50

34

188

Canada

95

25

53

39

212

Latin America

184

126

97

51

458

Latin America

206

140

116

68

530

Europe

458

312

242

141

1,153

Europe

486

360

258

157

1,261

Asia/Pacific

350

642

267

166

1,425

Asia/Pacific

376

700

291

184

1,551

Africa Middle East Total

58

40

65

29

192

Africa

193

366

221

162

942

Middle East

1,564

1,676

1,127

662

5,029

Total

65

51

64

35

215

210

389

230

161

990

1,692

1,837

1,196

748

5,473

Hydrocarbon Processing is published monthly. Second class postage paid at Houston, Texas.

Copyright © 2009 by Gulf Publishing Co. All rights reserved.

Permission is granted by the copyright owner to libraries and others registered with the Copyright Clearance Center (CCC) to photocopy any articles herein for the base fee of $3 per copy per page. Payment should be sent directly to the CCC, 21 Congress St., Salem, Mass.

01970. Copying for other than personal or internal reference use without express permission is prohibited. Requests for special permission or bulk orders should be addressed to the Editor. ISSN 0018-8190/01.

Special Computer Sorts

You can purchase a computer disk of a program, customized to suit your needs, from the HPI Construction Boxscore’s massive 36-year database (“history”) of projects. Your sort can be as narrow as the history of a particular type of Subscription price (includes both print and online versions): United States and Canada, one year project or as broad as a compilation of the entire 36-year $140, two years $230, three years $315. Airmail rate outside North America $175 additional a database. For a prompt quotation, send a description of the special sort desired to the Boxscore Coordinator at our year. Single copies $20, prepaid.

Houston office.

Publisher’s Note: Boxscore information is authenticated three times per year by confidential letter survey from the editor to operating companies, licensors, engineering companies and constructors. To be included in the survey, please contact Boxscore Coordinator via e-mail at boxscore@gulfpub.com. The publisher makes every effort to verify information but offers no guarantee as to current status of projects.

See page B-8 for licensor, engineering and construction companies’ abbreviations.

Gulf Publishing Company P.O. Box 2608, Houston TX 77252-2608 Phone: 713-525-4626, Fax: 713-525-4695 e-mail: boxscore@gulfpub.com

I

JUNE 2009 HYDROCARBON PROCESSING B-1

HPI Construction Boxscore Update Company

Plant Site

Project

Capacity Est. Cost Status Licensor

Engineering

Constructor

UNITED STATES Alaska Alaska

Tesoro Corp Denali

Kenai North Slope

Arizona

El Paso Corp

Eloy

California California California California California California Colorado

Valero Refining Co Tesoro Corp Shell Intl Prod Paramount Petr Corp Alon/Fina Oil & Chemical Alon/Fina Oil & Chemical Enterprise Products

Benicia Los Angeles Martinez Paramount Paramount Paramount Meeker

Hawaii

Tesoro Corp

Illinois Illinois Illinois Illinois Illinois

Benzene Reduction * Gas Treating

EX

None None

Storage, Natural Gas

90

3.5 Bcf 75 Mbpd None None 25 Mbpd 17 Mbpd 25 Mbpd 1.5 Bcfd

E 2012 E

Mustang Fluor

A 2010

* Scrubber Cogeneration * Crude Unit * Hydrocracker Iso Treating IsocrackingUnit Gas Processing (2)

TO

Kapolei

Benzene Reduction

RE

Secure Energy, Inc. Secure Energy, Inc. ConocoPhillips ConocoPhillips ConocoPhillips

Decatur Decatur Wood River Wood River Wood River

Gasifier Gasifier (2) Scrubber Scrubber (2) Scrubber (3)

500 500 60 60 30

Indiana

BP

Whiting

Coker, Delayed

102 Mbpsd

Kansas

Frontier El Dorado

El Dorado

Scrubber

40 Mbpd

Louisiana Louisiana

Placid Rfg Co Placid Rfg Co

Port Allen Port Allen

Hydrotreat, Gasoline Scrubber

0.3 MMgpd 30 Mbpd

Massachusetts Massachusetts

Hoegh LNG Hoegh LNG

Neptune terminal Neptune terminal

LNG Shuttle Regasification Vessel (1) LNG Shuttle Regasification Vessel (2)

Mississippi Mississippi Mississippi

Enerkem Technologies Ergon Refining Ergon Refining

Pontotoc Vicksburg Vicksburg

* Biofuel Plant Deasphalting, Propane Hydrotreater

New Mexico New Mexico New Mexico New Mexico New Mexico

Holly Corp Holly Corp Holly Corp Western Refining Navajo Rfg Co

Navajo Rfy Navajo Rfy Navajo Rfy Gallup Lovington

Distiller, Crude Hydrocrack, Gasoil Solvent Deasphalting (2) * Scrubber Crude Unit

North Dakota

Tesoro Corp

Mandan

Ohio Ohio

Husky Energy Inc Sunoco Inc

Lima Toledo

Oklahoma Oklahoma Oklahoma Oklahoma Oklahoma Oklahoma Oklahoma Oklahoma

Chesapeake Energy Midstream Energy ConocoPhillips Atlas Pipeline Atlas Pipeline Enogex Inc Terra Industries Inc Wynnewood Rfg Co

Pennsylvania Tennessee Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas Texas

Flint Hills Resources Shell Deer Park GreenHunter BioFuels GreenHunter BioFuels Energy Transfer PL Propylene LLC Atlas Pipeline DCP Midstream Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Motiva Enterprises LLC Ivanhoe Energy ConocoPhillips

Caldwell * Terminal, Petroleum Deer Park * Hydrogen Galena Park Biodiesel Galena Park Methanol Godley Cryogenic Gas Plant Houston Ship Channel Propylene WestTex Plant Cryogenic Gas Plant North Texas Cryogenic Gas Plant Port Arthur Cogeneration Port Arthur Coker, Delayed (2) Port Arthur Crude Unit Port Arthur Hydrocracker Port Arthur Hydrotreat, Naphtha Port Arthur Hydrotreater (2) Port Arthur Hydrotreater (3) Port Arthur Isomerization Port Arthur Reformer Port Arthur Sulfur Recovery Port Arthur Sulfur Recovery (2) Port Arthur Sulfur Recovery (3) Port Arthur Treater, Tail Gas Port Arthur Treater, Tail Gas (2) Port Arthur Treater, Tail Gas (3) San Antonio * Processing, Heavy Oil Sweeny Scrubber

Undisclosed Undisclosed Undisclosed

Chesapeake Energy Chesapeake Energy Chesapeake Energy

Undisclosed Undisclosed Undisclosed

Cryogenic Gas Plant Cryogenic Gas Plant (2) Cryogenic Gas Plant (3)

TO 200 MMscfd 60 MMscfd 60 MMscfd

14 7 7

U 2009 U 2009 U 2009

Thomas Russell Co. Thomas Russell Co. Thomas Russell Co.

Utah Utah

Tesoro Corp Tesoro Corp

Salt Lake City Salt Lake City

Benzene Reduction Hydrotreater

EX

None 8 Mbpd

55 25

E 2011 UOP C 2008 Axens

Mustang Mustang

Washington

Tesoro Corp

Anacortes

Benzene Reduction

EX

None

90

E 2011 UOP

Fluor

Hydrocrack, Resid Hydrotreater Hydrotreater Hydrotreater Hydrocracker (2) Upgrader

29 42 47 120 54 54

Mbpd Mbpd Mbpd Mbpd Mbpd Mbpd

U E U H H U

CLG CLG CLG CLG Fluor|CLG Fluor|Propak|Air Liquide

Hydrocrack, Resid * Gasifier Storage, Crude * Hydrocracker

47 500 BY 10 100

Mbpd MW Mbpd Mbpd

U 2010 CLG E 2015 Siemens C 2008 H CLG

EX

17

MW MW Mbpd Mbpd Mbpd

145 145

Mm3 Mm3

20 MMgpy 8 Mbpd 8 Mbpd

250

Belco Fluor S&B CLG CLG CLG Gas Liquids Eng

S&B

E 2012

Mustang

U U C U U

Siemens Energy Siemens Energy Belco Belco Belco

Siemens Energy Siemens Energy

U

FW

FW

C 2009 Belco

Belco

C 2008 Axens C 2008 Belco

Mustang Belco

2011 2011 2008 Belco 2010 Belco 2010 Belco

U U

2009 2010

Hamworthy Hamworthy

P U 2009 KBR U 2009

Wink Eng Wink Eng

BE&K BE&K

KP Engineering, LP KP Engineering, LP Lauren Belco KP Engineering, LP

KP Engineering, LP KP Engineering, LP Lauren

22

EX

None

32

E 2011

Mustang

* Hydrocracker Scrubber

RE

30 Mbpd 75 Mbpd

U 2012 CLG U 2009 Belco

CLG Belco

Grady County Panhandle Ponca City Roger Mills Co Sweetwater Undisclosed Woodward Wynnewood

Gas Treating Helium Extraction Scrubber Cryogenic Gas Plant Cryogenic Gas Plant (2) Cryogenic Plant * Urea * Hydrotreat, Gasoline

BY TO

C C C C C U E E

Gas Liquids Eng Thomas Russell Co. Belco Thomas Russell Co. Thomas Russell Co. Thomas Russell Co. UCSA KP Engineering, LP

UCSA KP Engineering, LP

ConocoPhillips

Trainer

* Alkylation, HF

RE

E 2012 UOP

S&B

ConocoPhillips

Valero Refining Co

Memphis

E 2010

KP Engineering, LP

KP Engineering, LP

TO

15 25 30 TO 125 EX 60 TO 120 480 13

LPG Recovery

MMcfd MMscfd Mbpd MMscfd MMscfd MMscfd m-tpd Mbpd

14 Mbpd

15 90 60

2010 Belco 2012 2009 CLG CLG CLG 2008

E C E H C

BY

40 15 18 11 70

E E U H H H C

Mbpd Mbpd Mbpd Mbpd Mbpd

Benzene Reduction

BY

None

300 43

10 8 14 6 14

220

20.7 MMscfd

RE 105 RE 1500 EX 200 544 TO 150 40 146 95 325 75 113 50 60 48 85 525 525 525 525 525 525 10

None None MMgpy bpd MMscfd MMmtpy MMscfd MMscfd MW Mbpd Mbpd Mbpd Mbpd Mbpd Mbpd Mbpd Mbpd tpd tpd tpd tpd tpd tpd bpd None

50 13 400 20 5

P P C C C U U C U U U U U U U U U E E E E E E C U

2009 2009 Process Dynamics Inc 2009 KBR Belco 2009

2008 2008 2009 Belco 2008 2008 2009 2010 UCSA 2010 Axens

2010 2010 2008 2007 2008 2010 2009 2008 2011 2011 2011 2011 2011 2011 2011 2011 2011 2009 2009 2009 2009 2009 2009 2009 2009

Linde

KP Engineering, LP

Linde

Burns and Roe ConocoPhillips Shell GSI CLG UOP Shell GSI CLG UOP UOP Black & Veatch Black & Veatch Black & Veatch Black & Veatch|Shell GSI Black & Veatch|Shell GSI Black & Veatch|Shell GSI

Thomas Russell Co. Lummus Technology Thomas Russell Co. Thomas Russell Co. Burns and Roe Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV|CLG Bechtel\Jacobs JV Bechtel\Jacobs JV

Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV

Bechtel\Jacobs JV Bechtel\Jacobs JV S&B S&B S&B S&B S&B S&B

Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV Bechtel\Jacobs JV

Belco

Belco

Lummus Technology

S&B

CANADA Alberta Alberta Alberta Alberta Alberta Alberta

North West Upgrading Fort Hills Energy Shell Canada Fort Hills Energy OPTI Canada Inc OPTI Canada Inc

Edmonton Edmonton Scotford Fort Hills Long Lake Long Lake

Alberta Alberta Alberta Alberta

Shell Canada EPCOR Power L.P. Devon Canada Corp Fort Hills Energy

Ft Saskatchewan Genesee Manatokan Sturgeon Lake

B-2

I JUNE 2009 HYDROCARBON PROCESSING

2010 2011 2010 2011 2011 2008

CLG CLG CLG CLG CLG

Bechtel|Bantrel|CLG Siemens Gas Liquids Eng CLG

KBR

Fluor|Ledcor| Air Liquide|Flint PCL

See page B-8 for licensor, engineering and construction companies’ abbreviations.

Company

Plant Site

Project

British Columbia British Columbia

Lignol Terasen Gas

Burnaby Mt. Hayes

* Biorefinery Storage, Natural Gas

Manitoba

Koch Chemical

Brandon

* Ammonia

New Brunswick New Brunswick New Brunswick

Irving Oil Ltd Irving Oil Ltd Irving Oil\BP JV

Eider Rock Eider Rock Eider Rock

* Hydrocracker * Hydrotreater Refinery

Saskatchewan Saskatchewan Saskatchewan Saskatchewan Saskatchewan Saskatchewan Saskatchewan Saskatchewan Saskatchewan Saskatchewan Saskatchewan Saskatchewan

Saskferco Products Inc Consumers Coop Refineries Consumers Coop Refineries Consumers Coop Refineries Consumers Coop Refineries Consumers Coop Refineries Consumers Coop Refineries Consumers Coop Refineries Consumers Coop Refineries Consumers Coop Refineries Consumers Coop Refineries Consumers Coop Refineries

Belle Plaine Regina Regina Regina Regina Regina Regina Regina Regina Regina Regina Regina

Urea (3) Cracker, FCC (2) Crude Unit Desulfurization, Diesel (2) Hydrotreat, Distillate (2) Hydrotreater, Naphtha Isomerization (2) Platformer (2) Polymerizer (1) Sat Gas (2) Storage, LPG Storage, Tank

Capacity Est. Cost Status Licensor 100 Ml/y 1.5 Bcfd

10 173

EX 1350 m-tpd 141 Mbpd 78 Mbpd 300 Mbpd RE 3400 m-tpd EX 22 Mbpd EX 30 Mbpd RE 8.5 Mbpd RE 18 Mbpd RE 6 Mbpd RE 6 Mbpd RE None RE None RE None 18 MMbbl 800 MMbbl

7000

1500

Engineering

Constructor

C 2009 U 2011

CB&I

CB&I

S 2012 ACSA

ACSA

ACSA

U 2013 CLG U 2013 CLG P 2015

CLG CLG

E E E E E E E E F E E E

2009 2012 2012 2011 2011 2011 2011 2011 2012 2011 2012 2012

Stamicarbon UOP UOP Axens UOP UOP UOP UOP UOP

Mustang|IAG Mustang|IAG Colt Eng Colt Eng Colt Eng Colt Eng Colt Eng Mustang Colt Eng Colt Eng Colt Eng

LATIN AMERICA Argentina Argentina Argentina

Repsol YPF Repsol YPF Repsol YPF

La Plata La Plata La Plata

* Coker, Delayed (replace) * Fractionator * Gas Plant

185 m3/hr None None

Brazil Brazil Brazil Brazil

Fosfertil-Ultrafertil Ultrafertil SA Petrobras Petrobras

* Ammonia Melamine Scrubber Lube Hydroprocessing

RE 1290 20 44 16

m-tpd Mtpy Mbpd Mbpd

E A E U

Brazil Brazil Brazil Brazil Brazil Brazil Brazil Brazil

Petrobas Petr Brasileiro SA Petrobras Quattor Petrobras RNEST Petrobras RNEST Petrobras RNEST Petrobras RNEST

Araucaria Araucaria REFAP Reduc (Duque de Caxias Refinery) Guanabara Bay Linhares Linhares Maua Pernambuco Pernambuco Pernambuco Pernambuco

LNG FSRU * Gas Treating * Processing, Oil * Cumene * Hydrogen (1) * Hydrogen (2) * Hydrotreater (1) * Hydrotreater (2)

138 EX 18 15 BY 110 125 125 83 83

Mm3 MMm3/d 200 Mbpd 200 Mm-tpy 40 MNm3/h MNm3/h Mbpsd Mbpsd

U 2009 U 2009

Chile

Enercon

Aconcagua Refinery

Amine Regeneration

Chile

Enercon

Aconcagua Refinery

Coker, Delayed

Chile Chile

Enercon Enercon

Aconcagua Refinery Aconcagua Refinery

Heater, Coker Stripper, Sour Water

Chile

Enercon

Aconcagua Refinery

Sulfur Recovery

45 tpd

C 2008

Chile Chile

Enercon Enercon

Aconcagua Refinery Aconcagua Refinery

Utilities Water Treatment

None 2280 m3/d

C 2008 C 2008

Chile

Undisclosed

Undisclosed

Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico Mexico

Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Petroleos Mexicanos Sonora Terminal and Pipeline

Burgos Cangrejera Cangrejera Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Minatitlan Morelos Morelos Morelos Puerto Libertad

Trinidad Trinidad

Phoenix Park Gas Processors Phoenix Park Gas Processors

Point Lisas Point Lisas

Venezuela Venezuela Venezuela Venezuela

Pequiven Pequiven Pequiven not disclosed

Jose Jose Puerto Nutrias Undisclosed

4135 m3/d

Butane Fractionator NGL Recovery * Urea (1) * Urea (2) * Urea * Gas Compression (2)

430

35 MW 1100 m3/d

EX EX

BY BY

FW FW FW

2010 ACSA 2003 2010 Belco 2015 CLG

2008 2011 2011 2011 2011

ACSA

Hamworthy UOP Haldor Topsøe Haldor Topsøe Haldor Topsøe Haldor Topsøe

C 2008 FW

H 2011 CLG

MMcfd Mtpy Mtpy Mtpd Mtpd Mtpd Mtpd Mtpd MMcfd Mtpd Mbpd tpd tpy Mtpy Mtpy Mtpy Bcfd

C E H U U U U U U U U U U H U U F

ACSA

CLG

C 2008 FW C 2008

45 Mbpd 200 210 100 13.4 13.4 55.8 42 150 48 37 50 7400 600 300 135 300 1

C E E E E

FW FW FW

C 2008

20 Mbpd

Hydrocracker * Cryogenic Gas Plant * Paraxylene * Styrene * Alkylation (1) * Alkylation (2) * Coker, Delayed * Cracker, FCC (2) * Distiller, Crude * Hydrogen * Hydrotreat, Distillate * Hydrotreat, Gasoil * Hydrotreat, Naphtha * Sulfur * Ethylene (2) * Ethylene Oxide (2) * Polyethylene (2) LNG Terminal

E E E

Promon Haldor Topsøe Haldor Topsøe Haldor Topsøe Haldor Topsøe

Platume

FW|MAN Ferrostaal| Tecnicas Reunidas MAN Ferrostaal| Tecnicas Reunidas|FW FW FW|Tecnicas Reunidas| MAN Ferrostaal FW|MAN Ferrostaal| Tecnicas Reunidas FW FW|MAN Ferrostaal| Tecnicas Reunidas CLG

Tecnicas Reunidas| MAN Ferrostaal|FW FW|Tecnicas Reunidas| MAN Ferrostaal FW Tecnicas Reunidas| MAN Ferrostaal|FW Tecnicas Reunidas| MAN Ferrostaal|FW FW Tecnicas Reunidas| MAN Ferrostaal|FW

2009 2011 2009 2009 2009 2009 2009 2009 2009 2009 2009 2009 2011 2009 CH.IV Intl

3500 bpd 600 MMscfd

C 2009 U 2009 Ortloff

Black & Veatch Black & Veatch

Black & Veatch Black & Veatch

2200 2200 2200 25

E E E P

Tecnimont Tecnimont Tecnimont Burckhardt Compression

Tecnimont Tecnimont Tecnimont

m-tpd m-tpd m-tpd MW

2012 Stamicarbon 2012 Stamicarbon 2013 Stamicarbon 2010

EUROPE Austria Austria

Neste Oil\OMV OMV AG

Schwechat Undisclosed

Biodiesel Gas Treating

200 Mtpy 3 MMNm3/d

Belarus Belarus

Grodno Azot Naftan Refinery

Grodno Novopolotsk

Belgium Belgium

ExxonMobil Chem Europe BASF\Dow Chemical Co. JV

Antwerp Antwerp

Bulgaria Bulgaria Bulgaria Bulgaria Bulgaria Bulgaria Bulgaria Bulgaria

Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas Lukoil Neftochim Bourgas

Burgas Burgas Burgas Burgas Burgas Burgas Burgas Burgas

Croatia Croatia Croatia

INA INA INA

Rijeka Sisak Sisak

Czech Republic

Sokolovska Uhelna, a.s.

Vresova

* Gasifier

200 MW

England England

SONHOE Dev. Co. SONHOE Dev. Co.

Teesside Wilton

Upgrader, Heavy Oil * Hydrocracker

200 Mbpd 200 Mbpd

Finland Finland Finland

Neste Jacobs Neste Oil Neste Oil

Porvoo Porvoo Porvoo

* FCC, flue gas Isomerization (1) Renewable Diesel (2)

None 600 Mm-tpy 170 Mm-tpy

France France

Total SolVin

Gonfreville Tavaux

Controls/Info Systems * Heater, Vacuum Cogeneration HPPO * Amine Recovery Hydrocrack, Gasoil Hydrocrack, Resid * Hydrogen (1) * Hydrogen (2) * Offsites * Sour Water Stripper * Utilities (2) Hydrocracker Diesel, HDS Hydrocracker

Lube Hydroprocessing Polyvinylidene Chloride

None None 125 MW 300 Mtpy 37 47 7500 7500

A 2008 Neste Jacobs H Black & Veatch

Black & Veatch

E 2008 U 2009

Siemens FW

C 2009 C 2009 BASF\Dow Chemical Co. JV

None Mbpd Mbpd kg/hr kg/hr None None None

F E E F F F F F

29 Mbpd 20 Mbpd 20 Mbpd 4000 4000 126

8 Mbpd 20 Mm-tpy

See page B-8 for licensor, engineering and construction companies’ abbreviations.

2012 2012 2012 2012 2012 2012 2012 2012

Axens Axens Axens Axens Axens Axens Axens Axens

Technip Technip Technip Technip Technip Technip Technip Technip

U 2010 CLG U 2013 CLG U 2013 CLG

CLG CLG CLG

C 2008 Siemens

Siemens

F 2014 A CLG

CLG

U 2011 Belco|SGS U 2010 Axens U 2009 Neste Oil

Belco Neste Jacobs Neste Jacobs

U 2011 CLG C 2009

CLG

Lummus Technology Lummus Technology Lummus Technology

Neste Jacobs Neste Jacobs

JUNE 2009 HYDROCARBON PROCESSING

I B-3

Company

Plant Site

Project

Capacity Est. Cost Status Licensor

France

Undisclosed

Undisclosed

Germany

Yara Brunsbuettel

Brunsbuettel

* Urea

Hungary Hungary Hungary

MOL Hungarian Oil & Gas MOL Hungarian Oil & Gas MOL Hungarian Oil & Gas

Danube Refinery Danube Refinery Szazhalombatta

* Hydrocracker Hydrotreat, Diesel Hydrocracker

Ireland

ConocoPhillips

Whitegate

* Sulfur Recovery

Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy Italy

Raffineria di Gela SpA Eni SpA Eni SpA Eni SpA Eni SpA Raffineria di Milazzo Eni SpA Eni SpA Eni SpA Eni SpA Eni SpA Eni SpA Eni SpA Eni SpA Eni SpA Eni SpA Eni SpA Eni SpA Eni SpA Eni SpA

Gela * Hydrogen Miglianico Field Sulfur Degasser Miglianico Field Sulfur Recovery Miglianico Field Sweetening, Gas Miglianico Field Treater, Tail Gas Milazzo Hydrodesulf (HDS) Raffineria di Sannazzaro * Sour Water Stripper Raffineria di Sannazzaro Sulfur Degasser Raffineria di Sannazzaro Sulfur Recovery Raffineria di Sannazzaro * Sulfur Recovery (1) Raffineria di Sannazzaro * Sulfur Recovery (2) Raffineria di Sannazzaro Treater, Tail Gas Raffineria di Sannazzaro * Treater, Tail Gas (2) Sannazzaro Hydrocracker Raffineria di Taranto Hydrogen Raffineria di Taranto Sulfur Degasser (3) Raffineria di Taranto Sulfur Recovery Raffineria di Taranto Treater, Tail Gas Taranto Hydrocracker Venice * Hydrocracker

Kazakhstan Kazakhstan Kazakhstan Kazakhstan Kazakhstan Kazakhstan Kazakhstan

LyondellBasell/SAT&Co LyondellBasell/SAT&Co LyondellBasell/SAT&Co LyondellBasell/SAT&Co LyondellBasell/SAT&Co LyondellBasell/SAT&Co LyondellBasell/SAT&Co

Atyrau Atyrau Atyrau Atyrau Atyrau Atyrau Atyrau

* Complex * Dehydrogenation, Propane * Ethane Cracker * Gas Separation * Polyethylene (1) * Polyethylene (2) * Polypropylene

Lithuania Lithuania Lithuania Lithuania

Mazeikiu Nafta Mazeikiu Nafta Mazeikiu Nafta Mazeikiu Nafta

Juodeikiai Mazeikiai Mazeikiai Mazeikiai

* Hydrocracker Alkylation, Solid-Acid Amine Recovery Utilities

Hydrogen

2.5 MNm3/h

Netherlands

NAM\Shell

Assen

Gate Terminal BV

Maasvlakte

* LNG Terminal (2)

Norway Norway Norway

TCM TCM StatoilHydro

Karsto Karsto Kollsness

* Amine Carbon Dioxide Capture * Gas Plant

Poland Poland

Anwil SA Anwil SA

Wloclawek Wloclawek

* Nitrogen Oxide Reduction Sys * Utilities

Portugal Portugal Portugal Portugal Portugal Portugal

Galp Energia Galp Energia Galp Energia Galp Energia Galp Energia Galp Energia

Porto Porto Porto Porto Porto Sines

Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation

Rosneft Kirishinefteorgsyntez Rosneft ZAO NaftaTrans ZAO NaftaTrans ZAO NaftaTrans ZAO NaftaTrans ZAO NaftaTrans ZAO NaftaTrans ZAO NaftaTrans Naftatrans ZAO NaftaTrans ZAO NaftaTrans ZAO NaftaTrans ZAO NaftaTrans ZAO NaftaTrans Korimos Nizhnekamskneftekhim TANECO TANECO Gazprom Gazprom Neft Sibur Khimprom Sakhalin Energy Investment Co Sakhalin Energy Investment Co Sibur Khimprom Sibur Khimprom Togliattiazot Rosneft Novomoskovsk Azot Novomoskovsk Azot Lukoil-Volgograd Neftepererabotk

Achinsk Kirishi Komsomolsk Krasnodar Krasnodar Krasnodar Krasnodar Krasnodar Krasnodar Krasnodar Krasnodar Krasnodar Krasnodar Krasnodar Krasnodar Krasnodar Moscow Nizhnekamsk Nizhnekamsk Nizhnekamsk Novy Urengoy Omsk Perm Sakhalin Island Sakhalin Island Tobolsk Tobolsk Togliatti Tuapse Tulskaya Oblast Tulskaya Oblast Volgograd

Serbia Serbia

NIS Pancevo Oil Refinery NIS Pancevo Oil Refinery

Pancevo Pancevo

Slovakia

Slovnaft as

Bratislava

* Polyethylene, LD

Spain Spain Spain Spain Spain Spain Spain Spain Spain Spain Spain Spain Spain Spain Spain Spain Spain

Repsol YPF Repsol YPF Repsol YPF Enagas Repsol YPF BP Oil Rfy de Castellon BP Oil Rfy de Castellon Enagas CEPSA REGANOSA Petronor Petronor Petronor Petronor Petronor Petronor Petronor

Cartagena Cartagena Cartagena Cartagena Cartagena Castellon Castellon Huelva La Rabida Refinery La Coruna Muskiz Muskiz Muskiz Muskiz Muskiz Muskiz Somorrostro

Coker, Delayed Heater, Coker Heater, Vacuum * LNG Terminal (4) Vacuum Unit Coker, Delayed Heater, Coker * LNG Terminal (3) Heater, Crude * LNG Terminal (2) * Butadiene * Cogeneration * Coker * Coker, Naphtha * Merox * Sulfur Coker, Delayed

B-4

I JUNE 2009 HYDROCARBON PROCESSING

Caloric

E 2010 UCSA

UCSA

26 Mbpd 4 Mbpd 36 Mbpd

U 2012 CLG U 2011 CLG C 2008 CLG

CLG

10 tpd

H

120 15 15 9 15 RE 62 3.5 210 160 80 80 275 160 27 120 210 160 340 13 21

Netherlands

Cond. Sweetening

E H H H H C F C C F F C F U U U U U U H

2011 2009 2009 2009 2009 2009 2009 2009 2009 2009 2009 2009 2009 2009 2009 2009 2009 2009 2009

None None None None None None None

P P P P P P P

2014 2014 2014 2014 2014 2014 2014

Siirtec Nigi Albemarle Siirtec Nigi Siirtec Nigi Siirtec Nigi Siirtec Nigi Siirtec Nigi Siirtec Nigi Siirtec Nigi CLG Haldor Topsøe Siirtec Nigi Siirtec Nigi Siirtec Nigi CLG CLG

Techint Siirtec Nigi|Irem Irem|Siirtec Nigi Irem|Siirtec Nigi Siirtec Nigi|Irem Techint SpA

Siirtec Nigi Siirtec Nigi

Siirtec Nigi Siirtec Nigi

Siirtec Nigi

Siirtec Nigi

Snamprogetti|CLG Techint Siirtec Nigi Siirtec Nigi Siirtec Nigi Snamprogetti|CLG CLG

TO 1510 bpd

U 2009 Merichem

Technisch Bureau Dahlman

E 2011

Techint|ENTREPOSE

E 2011 Aker Clean Carbon E 2011 Aker Solutions E 2011

Aker Clean Carbon Aker Clean Carbon Aker Solutions

E 2007 Chemeko E 2008 ILF Consulting Engineering

Chemeko Chemeko

E E E E E U

2011 2011 2011 2011 2011 2011 CLG

Fluor Fluor Fluor Fluor Fluor CLG

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

2013 2010 2012 2008 2008 2008 2008 2008 2008 2008 2008 2008 2008 2008 2008 2008

CLG CLG CLG Bechtel Bechtel Bechtel Bechtel Bechtel Bechtel Bechtel Bechtel|Vnipineft Bechtel Bechtel Bechtel Bechtel Bechtel

None None None 50 Mbpd None 2.5 MMtpy None 43 Mbpd

455 455

35 Mbpd 60 Mbpd 35 Mbpd None None None None None None None 6 MMtpy None None None None None RE 400 m-tpd EX 650 Mtpy 55 Mbpd 5 None 420 Mtpy 17 MMtpy 80 Mtpy 4.8 MMtpy 4.8 MMtpy 550 Mtpy 550 Mtpy RE 2600 m-tpd 82 Mbpd 1400 m-tpd 2000 m-tpd 6 MMtpy 30 Mbpd 30 Mbpd 220 Mtpy 53 Mbpd None None EX None 90 Mbpd 20 Mbpd 43.5 MW EX None None EX None 5 m-tpd 42 MW 2115 m-tpd 28 m-tpd 2115 m-tpd 110 m-tpd 36 Mbpd

1081

2009 2011 2012 2010 2010 2009 2008 2008 2010 2010 2010 2012 2010 2009 2009

CLG CLG CLG Shell GSI Shell GSI Shell GSI UOP Shell GSI UOP UOP UOP Shell GSI Shell GSI Shell GSI Shell GSI Exelus Lummus Technology CLG CLG Basell Shell Shell UOP INEOS UCSA CLG Stamicarbon Stamicarbon

CB&I Lummus|Vnipineft CLG CLG Tecnimont|Vnipineft Vnipineft|Honeywell Intl Vnipineft Chiyoda|Toyo Japan Chiyoda|Toyo Japan Tecnimont|Vnipineft Linde|Vnipineft UCSA CLG Chemoprojekt Chemoprojekt Vnipineft

U 2012 CLG U 2012 CLG

CLG MECS|CLG

F 2012

Tecnimont

U U U U U C C U U E E E E E E E U

2011 2011 2011 2011 2011 2008 2008 2011 2010 2011 2011 2011 2011 2011 2011 2011 2010

Techint Siirtec Nigi Siirtec Nigi Siirtec Nigi

LyondellBasell|KPI LyondellBasell|KPI LyondellBasell|KPI LyondellBasell|KPI LyondellBasell|KPI LyondellBasell|KPI LyondellBasell|KPI CLG Exelus FW FW

None 100 Mtpy None

UCSA

Techint Siirtec Nigi Siirtec Nigi Siirtec Nigi Siirtec Nigi Techint SpA

E 2012 CLG H Exelus A A

RE RE

* Hydrotreat, Diesel * Hydrocracker Hydrotreat, Diesel Cracker, Visbreaker Distillate, HDS Distiller, Crude Vac Gas Plant Hydrocracker Hydrotreat, Naphtha Isomerization Refinery Reformer CCR Sour Water Stripper Sulfur Recovery Treater, Amine Treater, Tail Gas * Alkylation Ethylene Hydrocrack, Coker Gas oil Lube Hydroprocessing Polyethylene Refinery Ethylene Complex LNG LNG (2) Dehydrogenation Polypropylene * Urea (3) Hydrotreat, Diesel Urea Granulation Urea Granulation (2) Distillation, Crude

Haldor Topsøe Siirtec Nigi Siirtec Nigi

35 Mbpd RE 3500 bpd RE None RE None

RE

Constructor

Jacobs Nederland BV

m-tpd tpd tpd MMcfd tpd Mbpd Mcfd tpd tpd tpd tpd tpd tpd Mbpd tpd tpd tpd tpd Mbpd Mbpd

12000 MMm3/y

Cracker, Visbreaker Diesel, HDS Distiller, Vac Refinery Utilities Hydrocracker

Diesel, HDS Hydrocracker

U 2009

RE 2000 m-tpd

RE

Engineering

FW

Technisch Bureau Dahlman ENTREPOSE|Techint

Remwil

Tecnicas Reunidas Neftechimproekt SNKP

Toyo Japan|Chiyoda Toyo Japan|Chiyoda UCSA Chemoprojekt Chemoprojekt

FW FW FW FW

FW

Axens FW Axens UOP Centry FW

FW FW FW

FW FW FW

FW FW Intecsa-Uhde|FW Sener FW Intecsa-Uhde|FW FW|Intecsa-Uhde Sener Sener|FW

See page B-8 for licensor, engineering and construction companies’ abbreviations.

Company

Plant Site

Project

Capacity Est. Cost Status Licensor

Ukraine Ukraine Ukraine Ukraine Ukraine

Kherson Oil Refinery Ukrtatnafta JSC Ukrtatnafta JSC Ukrtatnafta JSC Ukrtatnafta JSC

Kherson Kremenchug Kremenchug Kremenchug Kremenchug

Hydrocracker * Deisopentanizer TO * Hydrodesulf (HDS) TO * Naphtha HDT TO * Technology Consultancy ServicesTO

United Kingdom

Powerfuel Plc

Hatfield Colliery

* Treater, Adsorption

19 88 610 380

Mbpd None Mm-tpy Mm-tpy None

A P P P P

2010 2011 2011 2011 2011

CLG Axens

None

P 2013 UOP

Axens Axens

Engineering

Constructor

Ukrneftekhimproekt Ukrneftekhimproekt Ukrneftekhimproekt Ukrneftekhimproekt

ASIA/PACIFIC Australia Australia Australia Australia Australia Australia Australia

QGC QGC Santos Shell\Gladstone Ports Corp JV Woodside Energy Ltd Woodside Energy Ltd Woodside Energy Ltd

Curtis Island Curtis Island Gladstone Gladstone Karratha Karratha Karratha

* LNG * LNG (2) LNG * LNG Acid Gas Removal Fractionator LNG (5)

Australia Australia Australia

Woodside Energy Ltd EnCana Corp Nexus Energy

Karratha (Pluto LNG) Latrobe Valley Melbourne

Waste heat recovery unit (2) * Gasifier * FPSO

China China China China China

Dalian Petrochem WEPEC Erdos Union Chemical Co JianFeng Chemicals Guizhou Crystal Organic Chem

Dalian Dalian Erdos Fuling Guizhou

Hydrotreat, Resid * Scrubber, FCC Urea Urea Acetic Acid (2)

3000 52 TO 3520 2700 36

China China China China China China China China China China China China China China China China China China China China China China China

Giuzhou Jinchi CNOOC Oil & Petrochem Hulunbeier New Gold Shanxi Lanhua Chemical Shanxi Lanhua Chemical Shanxi Lanhua Chemical PetroChina Sinochem Shaanxi Carbonification Bei Yuan Chemical Shijiazhuang Chem & Fbr PetroChina Tianjin Kaiwei Group Shanghai Coking & Chem Sinopec\SK Energy JV Sinopec\SK Energy JV Dragon Aromatics Co Sinopec Shenhua Ningxia Coal Shenhua Ningxia Coal Shenhua Ningxia Coal Shenhua Ningxia Coal Shenhua Ningxia Coal

Guizhou Huizhou Hulunbeier Jincheng Jincheng Jincheng Lanzhou Quanzhou Shaanxi Shenmu Shijiazhuang Sichuan Tianjin Undisclosed Wuhan Wuhan Xiamen Yanshan Yinchuan Yinchuan Yinchuan Yinchuan Yinchuan

* Urea * Lube Hydroprocessing * Ammonia Gasifier Gasifier (2) Syngas Scrubber Hydrotreat, Resid * Methanol * Polyvinyl Chloride (PVC) * Ammonia * Hydrotreat, Resid * Lube Hydroprocessing * Wet Sulfuric Acid (WSA) Polyethylene, MD/HD Polypropylene Hydrocracker Scrubber Gasifier Gasifier (2) Gasifier (3) Gasifier (4) Gasifier (5)

1750 9 EX 1632 500 500 100 25 48 EX 2000 50 EX 700 48 2 100 300 200 70

India India

Gujarat Narmada Valley Fertilizer Gujarat Narmada Valley Fertilizer

Bharuch Bharuch

Air Separation Cogeneration

EX

3.8 MMtpy 3.8 MMtpy 4 MMtpy None None None 4.2 MMtpy

7700

None 500 MW None

500 500 500 500 500

F F C S C C C

2013 ConocoPhillips Ltd 2013 ConocoPhillips Ltd 2008

Bechtel Bechtel Bechtel|FW

Bechtel Bechtel

2008 2008 2008

FW|WorleyParsons Ltd FW|WorleyParsons Ltd WorleyParsons Ltd|FW

WorleyParsons Ltd|FW FW|WorleyParsons Ltd WorleyParsons Ltd| CB&I|FW FW

U 2010 P 2013 Siemens P

FW Siemens SBM

Mtpy None m-tpd m-tpd Mtpy

C E C U U

2008 2011 2008 2010 2010

CLG Belco Stamicarbon Stamicarbon Chiyoda

LPEC|CLG

m-tpd Mbpd m-tpd MW MW Mm3/y Mbpd Mbpd m-tpd Mtpy m-tpd Mbpd Mbpd tpd kty Mtpy Mbpd None MW MW MW MW MW

E U E E E E U U E U E U U E U U U H U U U U U

2009 2010 2010 2012 2012 2010 2009 2010 2010 2010 2010 2013 2012

UCSA CLG ACSA

UCSA CLG ACSA Siemens Energy Siemens Energy

6 MWh 33 MW

38 38

12 34

Belco CLG MCSA Chisso ACSA CLG CLG Haldor Topsøe 2011 INEOS 2011 2011 CLG Belco 2010 2010 2010 2010 2010

C 2005 Linde U 2010

CNCEC|Chengda Eng CECC Design Institute Chiyoda

CLG MCSA ACSA CLG CLG Haldor Topsøe

Dalian Chengda Eng |CNCEC Guizhou Crystal Organic UCSA ACSA Siemens Energy Siemens Energy Siemens Energy MCSA ACSA

SEI|CLG Siemens Energy Siemens Energy Siemens Energy Siemens Energy Siemens Energy

Siemens Energy Siemens Energy Siemens Energy Siemens Energy Siemens Energy

Linde BHEL

DIRECT FIRED HEATERS SINGLE SOURCE ENGINEERING AND FABRICATION

1640 S. 101st E. Avenue · Tulsa, OK 74128

Our comprehensive experience in heat transfer technology and related engineering disciplines ensures that we provide the process industry with state-of-the-art designs and manufacturing. Each system is custom designed for your project specific needs. We work closely with you to optimize the interrelation of thermal, mechanical and structural as well as instrumentation and control engineering disciplines. 3-D modeling of all components is done to prove dimensional accuracy for proper field fit up. Our sister companies, Express Metal Fabricators and St. George Steel, perform the fabrication for all North American projects.

Please forward your requests and inquiries to sales@expresstechtulsa.com or call (918) 622-1420.

www.expresstechtulsa.com Select 154 at www.HydrocarbonProcessing.com/RS JUNE 2009 HYDROCARBON PROCESSING

I B-5

Company

Plant Site

Project

Capacity Est. Cost Status Licensor

India India India India India India India India India India India India India India India India India India India India India India India India India India India India India India India India India India India

Gujarat Narmada Valley Fertilizer Gujarat Narmada Valley Fertilizer Gujarat Narmada Valley Fertilizer Gujarat Narmada Valley Fertilizer Bharat Oman Refineries Bharat Oman Refineries Chennai Petroleum (CPCL) ONGC Ltd Chambal Fertilizers & Chem Chambal Fertilizers & Chem Chambal Fertilizers & Chem Chambal Fertilizers & Chem Krishak Bharati Coop Ltd Krishak Bharati Coop Ltd Krishak Bharati Coop Ltd Krishak Bharati Coop Ltd Reliance Petr Ltd Reliance Petr Ltd Brahmaputra Cracker and Polymer Chennai Petroleum (CPCL) Chennai Petroleum (CPCL) Chennai Petroleum (CPCL) Chennai Petroleum (CPCL) Chennai Petroleum (CPCL) Mangalore Rfg & Petrochemicals Mangalore Rfg & Petrochemicals Indian Oil Corp Indian Oil Corp Ltd Indian Oil Corp Indian Oil Corp Ltd Indian Oil Corp Rashtriya Chemicals NOCL Indian Oil Corp Ltd HPCL

Bharuch Bharuch Bharuch Bharuch Bina Bina Chennai Dahej Gadepan Gadepan Gadepan Gadepan Hazira Hazira Hazira Hazira Jamnagar Jamnagar Lepetkata Manali Manali Manali Manali Manali Mangalore Mangalore Paradip Paradip Paradip Paradip Paradip Trombay Undisclosed Gujarat Refinery Vizag

Indonesia Indonesia

PT Patra SK PT Trans Pacific Petrochem

Dumai Tuban

Lube Hydroprocessing LPG Recovery

Japan

Mitsubishi Gas Chemical

Mizushima

Meta-xylene

Malaysia Malaysia Malaysia Malaysia Malaysia

Malaysia LNG Dua Petronas Penapisan Malaysian Rfg Co Malaysian Rfg Co Malaysian Rfg Co

Bintulu Melaka Melaka Melaka Melaka

LNG Lube Oil Refining Offsites (2) Refinery Utilities (2)

Pakistan

Fatima Fertilizer Co

Sadiqabad

Urea

Singapore Singapore

Lucite Intl ExxonMobil Chemical Asia Pacific

Jurong Jurong

MMA Petrochemicals

Singapore

Shell Eastern Petr

Pulau Bukom

South Korea South Korea South Korea South Korea South Korea

Hyundai Oilbank Co., Ltd. SK Energy Doosan Hyundai Oilbank Co., Ltd. PolyMirae

Daesan Incheon Undisclosed Undisclosed Yeosu

* Hydrotreat, Resid * Hydrocracker * IGCC * Scrubber, FCC * Polypropylene

66 Mbpd 40 Mbpd None 52 Mbpd None

Taiwan Taiwan Taiwan

Chinese Petroleum Corp Chinese Petroleum Corp Chinese Petroleum Corp

Kaohsiung Talin Tao Yuan

* Hydrocracker Lube Hydroprocessing Hydrotreat, Resid

12 Mbpd 6 Mbpd 70 Mbpd

H CLG U 2012 CLG H 2010 CLG

CLG CLG CLG

Thailand Thailand Thailand Thailand Thailand

Bangchak Petroleum PCL PTT Public Co Ltd Thai Oleochemicals Co Thai Oleochemicals Co Thai Oleochemicals Co

Bangkok Khanom Map Ta Phut Map Ta Phut Map Ta Phut

Reformer, Steam * Recovery, Condensate * Esters FAME Fatty Alcohols

25 15 200 100

C U U C C

CTCI|FW WorleyParsons WorleyParsons Uhde|WorleyParsons WorleyParsons |Uhde

Vietnam

Nghi Son Refinery

Nghi Son

* Hydrotreat, Resid

105 Mbpd

Algeria Algeria Algeria Algeria Algeria Algeria Algeria Algeria

Algeria Oman Algeria Oman Algeria Oman Algeria Oman Sorfert Algerie Sonatrach Sonatrach Sonatrach

Arzew Arzew Arzew Arzew Arzew El Merk Facility El Merk Facility Hassi Messaoud

Ammonia (1) Ammonia (2) * Urea (1) * Urea (2) Urea * Processing, Oil * Storage, NGL LPG Recovery

Egypt Egypt Egypt Egypt Egypt Egypt Egypt Egypt

Abu Qir Fertilizers Co Burullus Gas Co Burullus Gas Co Egyptian Methanex EAgrium EAgrium EHC Suez Oil Processing Co

Abu Qir Idku Idku Damietta Damietta Damietta Suez Suez

Urea (2) Gas Dehydration Gas Dehydration (2) Methanol Urea Granulation Urea Granulation (2) * Ammonium Nitrate * Ammonium Nitrate

Nigeria

Nigeria LNG Ltd

Bonny Island

Nigeria Nigeria Nigeria

Chevron Nigeria Ltd Chevron Nigeria Ltd NNPC

Escravos Izombe Warri

Repub S Africa Repub S Africa

ROMPCO ROMPCO

Komatipoort Komatipoort

Gas Compression Project Management Services

* Ethyl Acetate Methanol Nitric Acid (5) TDI Hydrocracker Hydrotreater * Hydrogen Generation Ethylene Ammonia (1) Ammonia (2) Urea (1) Urea (4) * Ammonia (5) * Ammonia (6) * Urea (1) * Urea (7) Sulfur Recovery (5) Sulfur Recovery (6) Polyethylene LLD/HD * Hydrocracker * Sulfur Degasser * Sulfur Recovery * Sulfur Recovery (2) * Treater, Tail Gas * Distillation, Crude * Distillation, VDU Paraxylene Petrochemicals Polypropylene Refinery Styrene Monomer Methanol (2) * Scrubber, FCC Treater, Tail Gas * Project Management Services

150 BY 30.6 150 150 35 36 21 1.1 RE 1900 RE 2000 RE 3500 RE 3300 RE 1890 RE 1890 RE 3325 RE 3325 675 675 220 RE 46 200 100 100 200 1.2 700 15 600 TO 220 600

tpd Mtpy tpd tpd Mbpd Mbpd Mtpd MMtpy m-tpd m-tpd m-tpd m-tpd m-tpd m-tpd m-tpd m-tpd tpd tpd Mtpy Mbpd tpd tpd tpd tpd None None MMtpy None Mtpy MMtpy Mtpy tpd None tpd None

4.35 11.15 329

981

100 100

10 Mbpd 480 m-tpd RE RE RE RE

2010 2006 2010 2010 2011 2011 2011 2012 2009 2009 2009 2009 2012 2012 2009 2009 2010 2013 2009 2009 2009 2009 2012 2012 2012 2012 2011 2009 2011 2009 2011

KBK Chem Plinke Chematur CLG CLG Haldor Topsøe Linde KBR KBR Snamprogetti KBR KBR Saipem Saipem Black & Veatch Black & Veatch INEOS CLG Siirtec Nigi Siirtec Nigi Siirtec Nigi Siirtec Nigi

Indian Oil UOP|Lummus Technology Haldor Topsøe Belco Black & Veatch Haldor Topsøe

Constructor

KBK Chem Staff

Staff

IBI Chematur EIL|CLG EIL|CLG Haldor Topsøe Samsung Eng|Linde PDIL PDIL PDIL PDIL KBR KBR PDIL|Saipem PDIL|Saipem Black & Veatch Black & Veatch

EIL EIL Samsung Eng

Reliance Reliance

CLG

Jacobs Jacobs FW FW FW FW CB&I Lummus|FW PDIL L&T|Toyo India PDIL

Rashtriya L&T

C 2008 CLG H 2008 Black & Veatch

CLG

70 Mm-tpy

U 2009 UOP

Chiyoda

Chiyoda

1.3 MMtpy 16 Mbpd None 45 Mbpd None

C C U U U

FW OGP|FW FW FW FW

FW

U 2009 Stamicarbon

Sojitz Corp|Kawasaki Plant Systems

Kawasaki Plant Systems|Sojitz Corp

C 2008 Lucite Intl U 2011

FW FW Led JV|WorleyParsons

U 2010

Chiyoda Singapore Pte

FW FW Led JV| WorleyParsons Chiyoda Singapore Pte

U U F E E

CLG CLG FW

1500 m-tpd 120 Mtpy None

Cracker, FCC-Resid

U C U U U U U E U U U U E E E E U U E E F F F F E E E E E E E U E E U

Engineering

150

TO 37400 bpd

None MMscfd Mtpy Mtpy Mtpy

19 16

2009 2007 EMRE 2009 2009 2009

2011 CLG 2012 CLG 2014 2011 Belco LyondellBasell

2008 2009 2009 2008 2008

UOP Uhde Inventa-Fischer AT Agrar-Technik|Cognis Uhde|AT Agrar-Technik|Cognis

U 2013 CLG

CLG

m-tpd m-tpd m-tpd m-tpd m-tpd Mbpd MMscfd MMscfd

E E E E U E E E

2010 2010 2010 2010 2010 2012 2012 2012

Haldor Topsøe Haldor Topsøe Snamprogetti Snamprogetti Stamicarbon

PDIL PDIL PDIL PDIL Uhde Petrofac Petrofac Shaw E & C

RE 2500 m-tpd TO 450 MMcfd TO 450 MMcfd 3.6 Mm-tpd 2000 m-tpd 2000 m-tpd 1.06 Mm-tpd 85 Mm-tpy

E U U U E E F C

2010 2009 2009 2010 2010 2010

Stamicarbon Advantica Advantica Davy Process|JM Stamicarbon Stamicarbon Carbon Holdings

Uhde Siirtec Nigi Siirtec Nigi Techint Uhde Uhde KBR IE-SA

FW

SKEC

CTCI WorleyParsons WorleyParsons Uhde|WorleyParsons Uhde|WorleyParsons

AFRICA 2100 2100 3675 3675 3450 98 600 24

LNG (7)

8.5 MMtpy

Hydrocracker GTL * Alkylation, HF

C 2007

34 Mbpd 34 Mbpd RE 341 m-tpd

U 2010 CLG C 2007 Chevron |Sasol P Exelus

None None

Technip|FW|JGC|KBR| Chiyoda|Snamprogetti Snamprogetti|JGC|KBR|CLG FW

Uhde Saipem|Shaw E & C Uhde Siirtec Nigi Siirtec Nigi Techint Uhde Uhde

KBR|Snamprogetti|JGC

U 2009 U 2009

FW FW

FW FW

C C C C C C

2008 2008 2008 2008 2008 2008

FW FW FW FW FW FW

FW FW FW FW FW FW

E E E E U

2012 2011 2011 2011 2011

Namvaran|Tecnicas Reunidas Namvaran|Tecnicas Reunidas Namvaran|Tecnicas Reunidas Namvaran|Tecnicas Reunidas

Hampa Asphalt Toos Asphalt Toos Asfalt Toos Asphalt Toos

MIDDLE EAST Bahrain Bahrain Bahrain Bahrain Bahrain Bahrain

BAPCO BAPCO BAPCO BAPCO BAPCO BAPCO

Sitra Sitra Sitra Sitra Sitra Sitra

Diethanolamine Gas Treating Sour Water Stripper Sour Water Stripper-2 (2) Sulfur Recovery Treater, Tail Gas

Iran Iran Iran Iran Iran

Golestan Petrochemical Bandar Imam Petrochemical Bandar Imam Petrochemical Bandar Imam Petrochemical Bandar Imam Petrochemical

Agh-ghala Bandar Imam Bandar Imam Bandar Imam Bandar Imam

Urea Butane Ethane Hexanes Pentane

B-6

I JUNE 2009 HYDROCARBON PROCESSING

RE

None None None None 220 tpd None 3250 787 544 290 390

m-tpd m-tpy m-tpy m-tpy m-tpy

204 204 204 204

Stamicarbon Tecnicas Reunidas Tecnicas Reunidas Tecnicas Reunidas Tecnicas Reunidas

See page B-8 for licensor, engineering and construction companies’ abbreviations.

Company

Plant Site

Project Propane Hydrotreat, Resid Isomerization Naphtha HDT Reformer CCR Urea Granulation Methanol (2) Distillation, VDU Urea Urea Granulation Urea (2) Polyethylene, LD

Capacity Est. Cost Status Licensor

Iran Iran Iran Iran Iran Iran Iran Iran Iran Iran Iran Iran

Bandar Imam Petrochemical Esfahan Oil Refinery Co Esfahan Oil Refinery Co Esfahan Oil Refinery Co Esfahan Oil Refinery Co NPC/Golestan Petrochem Kharg Petrochemical Lavan Oil Refinery Lordegan Petrochemical NPC/Lordegan Petrochem Co Ghadir Urea/Ammonia Petro Co Amir Kabir Petrochemical

Iran Iran Iraq

Zanjan Petrochemical NPC/Zanjan Petrochem Co SCOP

Bandar Imam Esfahan Esfahan Esfahan Esfahan Golestan Kharg Island Lavan Island Lordegan Lordegan Bandar Assaluyeh Petrochemical Special Economic Zone Zanjan Zanjan Karbala

Kuwait Kuwait Kuwait Kuwait Kuwait

KPPC KNPC KNPC KNPC Kharafi National

Shuaiba Industrial Area Naphtha, Light Al-Zour * Utilities Mina Abdulla Hydrotreat, Resid (2) Mina Al Ahmadi Hydrotreat, Resid Undisclosed Gas Treating

Qatar Qatar Qatar Qatar Qatar Qatar Qatar

Qatar Petroleum Qatar Petroleum Qatar Petroleum Qatar Petroleum Qatar Petroleum Qatar Fuel Additives ExxonMobil

Al Shaheen Al Shaheen Al Shaheen Al Shaheen Al Shaheen Mesaieed Ras Laffan

Scrubber Sulfur Recovery Sulfur Recovery (2) Treater, Tail Gas Treater, Tail Gas (2) * Methanol * Gas Processing (1)

Qatar

ExxonMobil

Ras Laffan

* Gas Processing (2)

Qatar Qatar Qatar Qatar Qatar

Qatar Shell GTL Ltd Oryx GTL Ltd Qatargas 2 Qatargas 2 RasGas (3)

Ras Laffan Ras Laffan Ras Laffan Ras Laffan Ras Laffan

Gas Processing Hydrocracker LNG (4) LNG (5) LNG (6)

Qatar

not disclosed

Undisclosed

* Gas Compression (1)

Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia

Arabian Petrochemical SHARQ Sipchem Saudi Aramco Saudi Aramco Sipchem Saudi Aramco Saudi Aramco PETRORabigh PETRORabigh PETRORabigh Aramco SABIC SABIC Sipchem

Al Jubail Al Jubail Al Jubail Manifa Bay Manifa Bay Jubail Ind City Khurais Khursaniyah Rabigh Rabigh Rabigh Ras Tanura Undisclosed Undisclosed Undisclosed

* ABS 200 kty MEG 700 Mtpy Vinyl Acetate 300 Mtpy Processing, Heavy Oil 900 Mbpd Project Management Services None Project Management Services None NGL 70 Mbpd * Utilities (2) EX None Offsites None Petrochemicals None Refinery None Sulfur Recovery 200 tpd * Acrylonitrile 200 Mtpy * Carbon Fiber 3 Mtpy * Ethylene Vinyl Acetate 200 Mtpy

Urea Urea Granulation * Refinery

1029 80 27 62 32 3250 4430 50 3250 3250 3250 300

m-tpy Mbpd Mbpsd Mbpsd Mbpsd m-tpd m-tpy Mbpd m-tpd m-tpd m-tpd Mtpy

204 74 187 187 188 632 58

230

3250 m-tpd 3250 m-tpd 140 Mbpd

2011 2012 2010 2010 2009 2012 2010 2009 2012 2012 2009 2009

Tecnicas Reunidas Axens UOP Axens Axens Stamicarbon Davy Process|JM Stamicarbon Stamicarbon Stamicarbon Basell

E 2012 Stamicarbon E 2012 Stamicarbon F 2010

800 Mtpy None RE 150 Mbpsd RE 50 Mbpsd 700 MMcfd 650 650 650 650 3000 850

E E E E E E H U E E C C

100

None tpd tpd tpd tpd m-tpd MMscfd

U P P U E

2009 UOP 2012 CLG 2011 CLG 2009

E H H H H S F

2012 2013 2013 2013 2013 2010 2009

850 MMscfd

F 2009

800 34 7.8 7.8 7.8

U U C C U

MMscfd Mbpd MMtpy MMtpy MMtpy

4 4 200

25 MW

2010 2011 2009 2009 2009

Belco Black & Veatch Black & Veatch Black & Veatch Black & Veatch MCSA Ortloff|GAA|EMRE| Merichem |UOP Ortloff|GAA|EMRE| Merichem|UOP Merichem |Shell CLG APCI|UOP UOP|APCI Shell|Ortloff|APCI| Merichem|UOP

P 2010

3204 3204 810

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

Engineering Namvaran|Tecnicas Reunidas Namvaran Namvaran|HEC Namvaran|HEC Namvaran Hampa Namvaran Namvaran Hampa Chiyoda|Toyo Japan Namvaran|Daelim| Simon Carves Hampa Technip SKEC|Bechtel|Tecnimont

Constructor Asphalt Toos Dorriz Dorriz Dorriz Hampa IIND Mehvar Hampa Hampa Steam|PIDEC|Chiyoda Namvaran|Simon Carves|Daelim Hampa Hampa SKEC|Tecnimont

CLG Gas Liquids Eng Technip Technip Technip Technip MCSA Chiyoda

MCSA

Chiyoda Chiyoda|HHI CLG Chiyoda\TechnipJV Chiyoda\TechnipJV Chiyoda\TechnipJV

Chiyoda|HHI Chiyoda\TechnipJV Chiyoda\TechnipJV Chiyoda\TechnipJV| Dodsal

Burckhardt Compression Shaw

2008 SD 2008 DuPont|Eastman 2011 2011 2008 2009 2008 2008 2008 WorleyParsons

FW FW FW FW HDEC|FW Petrofac FW FW FW FW

Fluor

HDEC|FW FW

2013 2013 2013

Process. Performance.

SOLUTIONS. OnQuest is a leader in process plant engineering and combustion technologies for clients in the petroleum and petrochemical industries. Our expertise includes efficient, energy-saving designs for hydrocarbon processing plants, with specialties in ammonia, hydrogen, syngas, LNG, and ethanol plants.

We are also a world leader in direct-fired process heater technology and burner management systems, and have particular expertise in lump-sum, turnkey projects, refurbishments, and revamps.

To learn more, call Randy Kessler at (909) 451-0502.

• High-Performance Solutions – Our designs prioritize project efficiency, for shorter field schedules, reduced man-hours, and lower costs.

• World-Class Expertise – Our engineers and process experts have decades of experience in complex design and installation projects.

• Global Capability – With offices in California, Texas, and Calgary, Alberta, and representatives in South America, Europe and Asia, we serve clients worldwide.

180 East Arrow Hwy | San Dimas CA 91773 Tel. (909) 451-0500 | Fax (909) 451-0499

www.onquest-inc.com PROCESS & FURNACE TECHNOLOGIES Select 155 at www.HydrocarbonProcessing.com/RS JUNE 2009 HYDROCARBON PROCESSING

I B-7

Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia

Company

Plant Site

SABIC SABIC SABIC Sipchem SABIC Osos Petrochemicals Osos Petrochemicals Osos Petrochemicals Osos Petrochemicals

Undisclosed Undisclosed Undisclosed Undisclosed Undisclosed Yanbu Yanbu Yanbu Yanbu

United Arab Emirates EPCL United Arab Emirates EPCL United Arab Emirates GASCO United Arab Emirates GASCO United Arab Emirates EPCL United Arab Emirates EPCL United Arab Emirates EPCL United Arab Emirates Borouge United Arab Emirates GASCO United Arab Emirates Borouge United Arab Emirates BorougeII United Arab Emirates BorougeII United Arab Emirates Takreer United Arab Emirates Undisclosed

Project

Capacity Est. Cost Status Licensor

* MMA * PMMA * Polyacetyl resins * Polyvinyl Acetate * Sodium Cyanide Offsites PBT Project Management Services Utilities

Dubai Hydrotreater Dubai Reformer, Cat Habshan Project Management Services Habshan Gas Complex Gas Plant Jebel Ali Offsites Jebel Ali Refinery Jebel Ali Utilities Ruwais Ethane Cracker Ruwais NGL (3) Ruwais Offsites Ruwais Polyethylene Ruwais Polypropylene Ruwais * Scrubber, FCC Undisclosed * Gas Processing, Sour

250 30 50 125 40

Mtpy Mtpy Mtpy Mtpy Mtpy None 60 Mtpy None None

EX RE RE RE EX EX EX EX

3204 3204 3204 810 3204

70 Mbpd None None 350 MMscfd None 120 Mbpd None 1.4 MMtpy None None 540 Mtpy 400 Mtpy None 60 MMscfd

S S S S S C C C C

2013 2013 2013 2013 2013 2008 2008 2008 2008

U U U U U U U U U U U U E

2009 2008 2009 2009 2009 2009 2009 2010 2009 2010 2010 Borealis A/S 2010 Borealis A/S 2015 Belco 2009

Engineering

Constructor

FW FW FW FW

FW

FW FW FW FW FW FW FW FW FW FW|Tecnicas Reunidas Tecnimont|FW Tecnimont|FW

FW FW

FW FW FW Tecnicas Reunidas Tecnimont Tecnimont

Epic Energy

SYMBOLS AND ABBREVIATIONS PROJECT or CAPACITY

* By To Re

First appearance in tabulation Increment of capacity added Total capacity after construction Revamp, modernize or de-bottleneck—not reported whether increment increase or final capacity Expansion—not classified

Ex

ABBREVIATIONS

bpd Bcf Bcfd Bcmy cfd cmd gpm kg/hr kgmol/hr kl kl/hr kty lb/d lb/hr LTPD m3/hr

barrel per day billion cubic feet billion cubic feet per day billion cubic meter per year cubic feet per day cubic meters per day gallons per minute kilograms per hour kilogram-mole per hour kiloliter kiloliters per hour kilotons per year pounds per day pounds per hour long ton per day cubic meters per hour

Mbpd Mcfd Mcfh Mcfy Mgpd Mm3 Mm3/d Mm3/hr MMbpd MMcfd MMl/y MMlb/y MMpcd MMscfd MMNm3/h MMNm3/y Mmtpd MMmtpy MMtpy MMtpd mt Mt mtpd Mtpy Mtpd MW

thousand barrels per day thousand cubic feet per day thousand cubic feet per hour thousand cubic feet per year thousand gallons per day thousand cubic meters thousand cubic meters per day thousand cubic meters per hour million barrels per day million cubic feet per day million liters per year million pounds per year million pounds per calendar day million standard cubic feet per day million normal cubic meters per hour million normal cubic meter per year thousand metric tons per day million metric tons per year million tons per year million tons per day metric ton thousand tons metric tons per day thousand tons per year thousand tons per day megawatt

MWh Nm3/d Nm3/h Scfd Sm3/h tpd tph tpy

megawatt per hour normal cubic meter per day normal cubic meter per hour standard cubic feet per day standard cubic meters per hour tons per day tons per hour tons per year

EST. COST

Cost in millions of US dollars.

This cost includes other units at this site and is repeated with these other units—whether they appear here or completed in earlier listings.

STATUS A Abandoned C Completed—deleted from subsequent tabulations E Engineering F Feed H Hold M Maintenance P Planning U Under construction T Presumed complete S Study

ABBREVIATIONS FOR Licensor, Engineering or Constructor AA AAT ABL ACS ACSA AET APCI AKC Astra Basell BDI BHEL BNI BPEC Burns Roe CB&I CCC CEPSA CLG CNCC CNCEC CPECC CWCEC Dedini ECC Ecolaire EDL EIED EIL EMRE Energea ENPPI

B-8

Azar Aab Ind. Co. Acid Amine Technologies Alliance Bechtel Linde Actividades de Construccion Servicios, S. A. Ammonia Casale SA Advanced Extraction Technologies Air Products & Chemicals, Inc. Aker Kvaerner China Astra Evangelista SA Basell Polyolefine GmbH (Sinopec) Beijing Design Institute Bharat Heavy Electricals Ltd. Ballast Nedam International B.V. (Sinopec) Beijing Petrochemical Eng. Burns and Roe Worley Chicago Bridge & Iron Co. N.V. Consolidated Contractors Co. Compañía Española de Petróleos Chevron Lummus Global China National Chemical Const.

China National Chemical Eng. Corp.

China Petroleum Eng & Constr Corp China Wuhuan Chem. Eng. Corp.

Dedini S/A Indústrias de Base Erection & Construction Co. Ecolaire Espana SA EDL Anlagenbau Gesellschaft MBH Energy Industries, Eng. & Design Engineers India Ltd. ExxonMobil Research & Engineering Energea Unwelttechnologie GmbH Engineering for the Petr & Process Ind

FB&D FFBL FRIPP FW GAA GE Genpro GLF GPN GV HDEC HEC HHI HQCEC HTI IAG ICI IET IFP IHI IMP IKPT IOEC IRSL ITL JGC Phil JJC JM JP KBC KBR

I JUNE 2009 HYDROCARBON PROCESSING

Ford Bacon & Davis Co., Inc. Fauji Fertilizer Bin Qasim Ltd. Fushun Research Institute of Petr & Petrochemicals Foster Wheeler Corp.

Goar Allison & Assoc.

General Electric Co. Genpro Engenharia Ltda Grand Lavori Fincosit SpA GPN Engineering & Process Giammarco-Vetrocoke Hyundai Engineering & Construction Hyundai Engineering Co. Ltd. Hyundai Heavy Industries Huanqiu Contracting & Eng. Corp.

Hydrocarbon Technologies, Inc. International Alliance Group Imperial Chemical Industries plc Integrated Environmental Tech Institut Français du Pétrole Ishikawajima Harima Heavy Ind. Instituto Mexicano de Petróleo PT Inti Karya Persada Tehnik Iranian Offshore Eng & Constr.

Indo Rama Synthetics Ltd Independent Technology Ltd. JGC Philippines JJC Contratistas Generales SA Johnson Matthey Catalysts Jahan Pars KBC Advanced Technologies plc Kellogg, Brown & Root, Inc.

KJT KPT KTY KVT Leighton LGC LGI LGNCH L&T LPEC MCC MCSA MEPI MGC MHI MES MWKL NIOEC Niplan NPCC NRC OIEC OGP OPD PCL PCT Engr PDE PDF PDIL PECL PIDEC

KBR/JGC/Technip JV Kvaerner Process Technology Consultoria e Projeto de Instalaçóes Industriais Kanzler Verfahrentechink GmbH Leighton Contractors Pty Ltd. LG Chemical Ltd. Le Gaz Integral OOO Lengironeftechim Larsen & Toubro Ltd. Luoyang Petrochemical Eng. Corp.

Mitsubishi Chemical Corp.

Methanol Casale SA Middle East Project Intl.

Mitsubishi Gas Chemical Mitsubishi Heavy Industries Ltd. Mitsui Engineering & Shipbuilding M W Kellogg Ltd. National Iranian Oil Eng. & Constr.

Niplan EngenhariaLtda National Petroleum Construction Co. Navajo Refining Co. Oil Industries Eng. and Constr.

OGP Technical Services Optimized Process Designs, Inc. PCL Industrial PCT Engineers Pty Ltd. Project Design Engineers Ltd. Process Design & Fabrication Pty Ltd. Projects and Development India Ltd. Pacific Engineers & Constructors Ltd. Petrochemical Ind. Design & Eng. Co.

PRAJ Projectus REG RIPP S&B SD SCO SEI Setal SIAC SKEC Shaw E & C SNEC Staff STC Takreer TBD TCE TJ UCC UCSA VEC VTA

PRAJ Industries Ltd. Projectus Consultoria Ltda Renewable Energy Group Research Institute of Petr Processing S&B Engineers & Constructors Scientific Design Co., Inc. Schrader-Camargo/Otepi Sinopec Engineering Inc. Setal Engenharia e Construcoes SA SIAC Butlers Steel Ltd. SK Engineering & Construction Co. Shaw Energy & Chemicals Sinopec Ningbo Eng. Co. Staff Service & Technology Corp.

Abu Dhabi Oil Refining Co To be determined TCE Consulting Engr., Ltd. Tehran Jonoob Union Carbide Corp.

Urea Casale SA VEC Ingenieria y Construccion Verfahrenstechnik & Automatisierung GmbH WEC Wuhuan Engineering Co. Ltd. WGII Washington Group International Inc. YFT Yara Fertilizer Technology Yara-KT Yara-Kaltenbach Thuring ZA Zoha Sonat ZRCC Design Zhenhai Refining & Chemical Co.

See page B-8 for licensor, engineering and construction companies’ abbreviations.

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Custom Article

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WANTED: MANUFACTURER’S REPS Dorf Ketal Chemicals seeks Manufacturer’s Reps to represent our process control chemicals in the Refinery market and the Ethylene sector. Ideal candidates will have significant experience and live near such plants. Contact: wloven@dorfketalusa.com

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Visit our Website at www.HydrocarbonProcessing.com HYDROCARBON PROCESSING JUNE 2009

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FREE Product and Service Information — JUNE 2009 HOW TO USE THE INDEX: The FIRST NUMBER after the company name is the page on which an This information must be proadvertisement appears. The SECOND NUMBER, appearing in parentheses, after the company name, vided to process your request: is the READER SERVICE NUMBER. There are several ways readers can obtain information: PRIMARY DIVISION OF INDUSTRY 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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This Advertisers’ Index and procedure for securing additional information is provided as a service to Hydrocarbon Processing advertisers and a convenience to our readers. Gulf Publishing Co. is not responsible for omissions or errors.

(check one only): A B C F G H J P

䊐-Refining Company 䊐-Petrochemical Co. 䊐-Gas Processing Co. 䊐-Equipment Manufacturer 䊐-Supply Company 䊐-Service Company 䊐-Chemical Co. 䊐-Engrg./Construction Co.

JOB FUNCTION (check one only): B E F G I J

䊐-Company Official, Manager 䊐-Engineer or Consultant 䊐-Supt. or Asst. 䊐-Foreman or Asst. 䊐-Chemist 䊐-Purchasing Agt.

ADVERTISERS in this issue of HYDROCARBON PROCESSING Company Website

Page

RS#

ABV Srl . . . . . . . . . . . . . . . . . . . . . . . . . 47

(177) (76) (110)

www.info.hotims.com/25253-110

Ashland Inc . . . . . . . . . . . . . . . . . . . . . . 37

(95)

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Axens . . . . . . . . . . . . . . . . . . . . . . . . . . 92

(53)

www.info.hotims.com/25253-53

Bently Pressurized Bearing Co . . . . . . . . 76

(173) (114) (113) (57) (156)

www.info.hotims.com/25253-156

Chemstations Inc. . . . . . . . . . . . . . . . . . 28

(153) (179) (71)

www.info.hotims.com/25253-71

Dresser-Rand. . . . . . . . . . . . . . . . . . . . . 74

(171)

www.info.hotims.com/25253-171

Dyna-Therm . . . . . . . . . . . . . . . . . . . . . 24

(152)

www.info.hotims.com/25253-152

Eaton . . . . . . . . . . . . . . . . . . . . . . . . . . 51

(116)

www.info.hotims.com/25253-154

GPC Software Video Books . . . . . . . . . 78 GPC Software Video Books . . . . . . . . . 84 Gulf Coast Turnaround Directory Showcase . . . . . . . . . . . . . . 69 Haver & Boecker . . . . . . . . . . . . . . . . . . 75

Inpro / Seal Company . . . . . . . . . . . . . . 10

(174) (170) (175)

www.info.hotims.com/25253-92

Peco Facet. . . . . . . . . . . . . . . . . . . . . . . 29 Prosim . . . . . . . . . . . . . . . . . . . . . . . . . 56

(168) (155) (98) (87) (164)

www.info.hotims.com/25253-164

Samson GmbH . . . . . . . . . . . . . . . . . . . . 4

(151)

(79) (167)

(66)

www.info.hotims.com/25253-66

Thermo Fisher Scientific . . . . . . . . . . . . . 52 (97)

(74)

www.info.hotims.com/25253-74

T.D. Williamson . . . . . . . . . . . . . . . . . . . 91 (96)

(161)

www.info.hotims.com/25253-161

Swagelok Co. . . . . . . . . . . . . . . . . . . . 6-7 (89)

(62)

www.info.hotims.com/25253-62

Sulzer Chemtech, USA Inc.. . . . . . . . . . . 43 (88)

(160)

www.info.hotims.com/25253-160

Spraying Systems Co . . . . . . . . . . . . . . . 32 (165)

(166)

www.info.hotims.com/25253-166

Soteica LLC . . . . . . . . . . . . . . . . . . . . . . 42 (51)

www.info.hotims.com/25253-97

Mangiarotti SpA . . . . . . . . . . . . . . . . . . 49

Paharpur Cooling Towers, Ltd. . . . . . . . . 18

Siirtec Nigi SpA . . . . . . . . . . . . . . . . . . . 62 (172)

(115)

www.info.hotims.com/25253-115

Tray-Tec Inc. . . . . . . . . . . . . . . . . . . . . . 46

(162)

www.info.hotims.com/25253-162

Turbomachinery Laboratory . . . . . . . . . . 44

(98)

www.info.hotims.com/25253-98

www.info.hotims.com/25253-167

(154)

Newton's . . . . . . . . . . . . . . . . . . . . . . . 66

www.info.hotims.com/25253-151

www.info.hotims.com/25253-96

M3 Technology . . . . . . . . . . . . . . . . . . . 66

(158)

www.info.hotims.com/25253-87

www.info.hotims.com/25253-79

(69)

Microtherm . . . . . . . . . . . . . . . . . . . . . . 38

www.info.hotims.com/25253-98

www.info.hotims.com/25253-89

KTI Corporation . . . . . . . . . . . . . . . . . . . 22

(99)

www.info.hotims.com/25253-155

(169)

www.info.hotims.com/25253-88

KBR . . . . . . . . . . . . . . . . . . . . . . . . . . . 26

MBI Leasing LLC . . . . . . . . . . . . . . . . . . 20

OnQuest . . . . . . . . . . . . . . . . . . . . . . . . 31

www.info.hotims.com/25253-51

HPI Marketplace . . . . . . . . . . . . . . . 86-88 Hytorc . . . . . . . . . . . . . . . . . . . . . . . . . . 60

RS#

www.info.hotims.com/25253-168

www.info.hotims.com/25253-172

Linde Process Plants . . . . . . . . . . . . . . . 12

www.info.hotims.com/25253-69

Express Integrated Technologies. . . . . . . 30

Events . . . . . . . . . . . . . . . . . . . . . . . . . 82

KTI Corporation . . . . . . . . . . . . . . . . . . . 25

www.info.hotims.com/25253-116

Emerson Process Management (Fisher Controls) . . . . . . . . . . . . . . . . . 16

(93)

www.info.hotims.com/25253-93

Gulf Publishing Company Circulation . . . . . . . . . . . . . . . . . . . . . 85 European Turnaround Directory . . . . . . 72 Events . . . . . . . . . . . . . . . . . . . . . . . . . 68

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www.info.hotims.com/25253-158

www.info.hotims.com/25253-165

www.info.hotims.com/25253-179

Costacurta SpA Vico . . . . . . . . . . . . . . . 64

Flexitallic LP . . . . . . . . . . . . . . . . . . . . . . 5

Honeywell International. . . . . . . . . . . . . . 2

www.info.hotims.com/25253-153

Coade Engineering Software . . . . . . . . . 70

(163)

www.info.hotims.com/25253-175

www.info.hotims.com/25253-57

Carver Pump Company . . . . . . . . . . . . . 34

Flexim GmbH . . . . . . . . . . . . . . . . . . . . 46

Company Website

www.info.hotims.com/25253-99

www.info.hotims.com/25253-170

www.info.hotims.com/25253-113

Carpenteria Corsi Srl . . . . . . . . . . . . . . . 39

(159)

www.info.hotims.com/25253-174

www.info.hotims.com/25253-114

Bryan Research & Engineering . . . . . . . . . 8

FBM Hudson Italiana SpA . . . . . . . . . . . 41

www.info.hotims.com/25253-169

www.info.hotims.com/25253-173

BME Global Limited. . . . . . . . . . . . . . . . 67

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www.info.hotims.com/25253-163

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Altair Strickland. . . . . . . . . . . . . . . . . . . 14

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www.info.hotims.com/25253-177

ACS Industries Inc. . . . . . . . . . . . . . . . . 63

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89

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

APC designs for minimum maintenance—Part 1 I, and others, have written much about the difficulties of maintaining advanced process control (APC) applications.1–5 It is a problem to be tackled on many levels. There are issues of manning, training and organizing that we APC engineers have only a limited influence on. On the other hand there are other areas that we do indeed influence, and this editorial addresses one of those: Design of the APC application in a way that would make it easily maintainable. What is more important for the APC—to recover 100% of potential benefits or to recover 70% of potential benefits and be easy to comprehend and maintain? Considering the history of APC maintenance the latter is better by far.

I would like to offer several rules for making APC more robust and durable, to the point that applications can survive in an environment of poor maintenance, but lack of space permits only one rule, leaving the others to be addressed by following editorials. The rule is generic and should apply to any APC configuration, though the simple distillation column of Fig. 1 will serve to illustrate the points.

Design rule 1. Do not clutter the control matrix.

Associate each control variable (CV) preferably with one, hopefully no more than two, manipulated variables (MVs).

But aren’t we dealing with a multivariable predictive control (MVPC) tool? Can’t we move many MVs to bring a CV to its desired target? Yes, the CV of interest could have a model against many MVs, but the easy-to-maintain (and to implement) application would move only one or two handles per CV. If those two MVs are off then the associated CV should be shed. In our distillation example, CV1 is a top product purity inference and MV1, the associated manipulated variable, is tray six temperature. Are there other MVs that could affect the top product purity? Yes, of course, increasing reboiler steam,

MV3 PC

MV2, would increase fractionation and affect the top purity. Increasing column pressure, MV3, would change the equilibrium on tray six and affect top product purity as well.

How then would the MVPC know that pressure and steam are not to be used to control top purity? How would the top purity be controlled when tray six TC is against a max limit?

Answer. Associate the column pressure, MV3, also with the top quality. Now there are two handles for control of top quality. Set one of these to a lower priority: Pressure changes are to take place only when the tray temperature controller is against a limit. In a well-operated application the tray temperature setpoint would not be superficially bound, but in a poor maintenance environment, if the operator makes a mistake of setting a temperature limit quality, control would still work.

And how should we use reboiler steam, MV2? The column of our example has a DCS tray temperature controller manipulating reflux. Increasing reboiler steam heats up the column, and the tray temperature controller would close the heat balance by increasing the reflux. This MV has only a limited influence on the top product quality, and I would be inclined to use it solely to control reboil ratio, CV3. A reasonable reboil ratio would ensure that if top purity specification is met then bottom purity is also under control.

To complete this design we ought to consider abnormal constraints, such as a hydraulic reflux valve limit, CV2, or a flooding limit, CV4. Should the demand for reboil ratio conflict with reflux valve position or flooding constraint that renders the control problem infeasible, dictating that the reboil ratio target, CV3, must be abandoned.

Thus, we have come up with the simplest possible design, and if the top product inferential model is reliable, such a design is likely to survive in a lack of maintenance environment. HP 1

TI

2

LC

FC

LC

TC

Tray 6

MV1

FC

3 4

FC

5

PI LC TI

FIG. 1

90

FC

Tray 25

6

TI

TI

Tray 30

A distillation column candidate for APC.

I JUNE 2009 HYDROCARBON PROCESSING

MV2 FC Steam

LITERATURE CITED Friedman, Y. Z., “Avoid advanced control project mistakes,” Hydrocarbon Processing, October 1992.

Latour, P. R., “Does the HPI do its CIM business right?,” guest editorial, Hydrocarbon Processing, July 1997.

Friedman, Y. Z., “Advanced process control‚it takes effort to make it work,” HPIn Control, Hydrocarbon Processing, February 1997.

Kane L. A., “Controversy in Control,” editorial, In Control, March/April 1998.

Friedman, Y. Z., “Audit your APC applications,” Hydrocarbon Processing, December 2006.

Friedman, Y. Z., “Choosing inferential modeling tools,” Hydrocarbon Processing, January 2006.

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.

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