HP_2012_03 by androsov.info

MARCH 2012

HPIMPACT

SPECIALREPORT

BONUSREPORT

BPâ&#x20AC;&#x2122;s new energy outlook

CORROSION CONTROL

CLEAN FUELS

New detection methods methods identify root causes for metal fatigue and failure

Process and product optimization yield improved fuels

Using zeolites for carbon capture

www.HydrocarbonProcessing.com

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SPECIAL REPORT: CORROSION CONTROL

Reduce CO2 in acid gas from amine-based TGTUs Improve furnace temperature and sulfur recovery B. Spooner and F. Derakhshan

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Consider a new monitoring system to prevent corrosion Innovative, continuous supervising method collects real-time data on key asset health P. Collins

Avoid stress corrosion cracking of stainless steel This case history investigates equipment failure in a glycol unit A. D. Jain

Operating philosophy can reduce overhead corrosion Boost refinery reliability by controlling potential amine recycle loops M. Dion, B. Payne and D. Grotewold

Closed-loop control can clamp down on crude unit corrosion Automating the detection process and controlling applications in real time dramatically improves performance N. P. Hilton

BONUS REPORT: CLEAN FUELS

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Cover KBR provided mechanical services for nearly 45 plant projects for Shell Canada Ltd. in Scotford, Alberta, Canada, including civil, structural steel erection, rail-car loading, sulfur degassing and insulation for the upgrader complex. Photo courtesy of KBR.

HPIMPACT 17

BP’s energy outlook to 2030

Octagonal opportunity for carbon capture

Choose a facility configuration based on financial metrics LP modeling can provide an unbiased, cost-based preview of a refinery/petrochemical plant design T. E. Swaty

Optimize hydrogen management for distillate production New tools enable refiners to fine-tune refinery configuration and maximize profits P. Parihar, R. Kumar, R. K. Voolapalli and S. Agarwal

GLOBAL TURNAROUND AND MAINTENANCE 2012—SUPPLEMENT

T-69

Executive leadership—Essential to ensure world-class turnarounds Many factors are involved in achieving ‘safe, reliable and cost-efficient’ turnarounds B. Singh

COLUMNS 11

HPINSIGHT Technology and market share: Challenges requiring ever-changing solutions

HPIN RELIABILITY Consider procedure cards, technical books and maintenance conferences for training

HPIN WATER MANAGEMENT Water is the next oil

HEAT TRANSFER DEVELOPMENTS

Achieve optimal heat recovery in a kettle exchanger Improve operations by avoiding buildup of sensible duty or by using baffles T. Das

PROCESS CONTROL AND INSTRUMENTATION

Advanced process control: A historical perspective A blend of art and science with a history worth recounting M. C. Delaney

ENGINEERING CASE HISTORIES

Case 67: Extruder blowback Understanding what is happening can lead to a solution T. Sofronas

DEPARTMENTS 9 HPIN BRIEF • 21 HPINCONSTRUCTION 28 HPI CONSTRUCTION BOXSCORE UPDATE 95 HPI MARKETPLACE • 97 ADVERTISER INDEX

HP ONLINE EXCLUSIVES HPINNOVATIONS

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Because Hydrocarbon Processing is edited specifically to be of greatest value to people working in this specialized business, subscriptions are restricted to those engaged in the hydrocarbon processing industry, or service and supply company personnel connected thereto. Hydrocarbon Processing is indexed by Applied Science & Technology Index, by Chemical Abstracts and by Engineering Index Inc. Microfilm copies available through University Microfilms, International, Ann Arbor, Mich. The full text of Hydrocarbon Processing is also available in electronic versions of the Business Periodicals Index.

Publisher Bill Wageneck Bill.Wageneck@GulfPub.com ARTICLE REPRINTS

If you would like to have a recent article reprinted for an upcoming conference or for use as a marketing tool, contact Foster Printing Company for a price quote. Articles are reprinted on quality stock with advertisements removed; options are available for covers and turnaround times. Our minimum order is a quantity of 100.

EDITORIAL Editor Stephany Romanow Reliability/Equipment Editor Heinz P. Bloch Process Editor Adrienne Blume Technical Editor Billy Thinnes Online Editor Ben DuBose Associate Editor Helen Meche Contributing Editor Loraine A. Huchler Contributing Editor William M. Goble Contributing Editor Y. Zak Friedman Contributing Editor ARC Advisory Group

For more information about article reprints, call Rhonda Brown with Foster Printing Company at +1 (866) 879-9144 ext 194 or e-mail rhondab@FosterPrinting.com. HYDROCARBON PROCESSING (ISSN 0018-8190) is published monthly by Gulf Publishing Company, 2 Greenway Plaza, Suite 1020, Houston, Texas 77046. Periodicals postage paid at Houston, Texas, and at additional mailing office. POSTMASTER: Send address changes to Hydrocarbon Processing, P.O. Box 2608, Houston, Texas 77252. Copyright © 2012 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.

MAGAZINE PRODUCTION Director—Production and Operations Sheryl Stone Manager—Editorial Production Angela Bathe Artist/Illustrator David Weeks Manager—Advertising Production Cheryl Willis

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Registration

NOW OPEN

MILAN, ITALY | 12–14 JUNE

Hydrocarbon Processing’s International Reﬁning and Petrochemical Conference (IRPC) is the leadingedge technical conference providing a forum within which you will be able to learn from and network with today’s industry leaders. This year’s event will give special focus to the areas of unconventional feedstocks and heavy oil conversion. Additional topics scheduled for discussion at the conference include, but are not limited to: • Heavy oil conversion/bottom-of-the-barrel

• Future of fuel oil

• Plant and reﬁnery sustainability

• Clean fuels

• Reﬁnery and petrochemical integration

• Maintenance and reliability

• Bio-based petrochemicals/chemicals

• Flare systems

• Alternative feedstocks-shale gas, GTLs, CTLs, etc.

• Process control applications/automation

To view the agenda and to learn more about IRPC 2012, please visit www.HPIRPC.com.

Register today to take advantage of Early Bird Pricing. Go to www.HPIRPC.com, or contact Gwen Hood at +1 (713) 520-4402 or Gwen.Hood@GulfPub.com

Early Bird Pricing

2012 IRPC Sponsors:

(before 30 April 2012)

Delegate Fees USD $930.00

Single Team of Two *

Pack of 10

Event Host Sponsor

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USD $1,674.00 USD $8,415.00

*Pack of 10 purchase includes a reserved table at lunch, listing as a Team Pack Sponsor in the event program, and signage with your company name and logo displayed at the conference.

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2012 International ReďŹ ning and Petrochemical Conference Advisory Board: Giacomo Rispoli Senior Vice President Research & Development IRPC Advisory Board Chair eniâ&#x20AC;&#x201C;Refining & Marketing Division John Baric Licensing Technology Manager Shell Global Solutions International B.V. Eric Benazzi Marketing Director Axens Carlos Cabrera Executive Co-Chairman Ivanhoe Energy Dr. Charles Cameron Head of Research & Technology BP Antonio Di Pasquale Vice President Refining Product Line Technip Giacomo Fossataro Technical and Operation Manager Walter Tosto S.p.A.

Mohammed Al-Gahtani Head of Refining Technical Group Saudi Aramco Dr. Madhukar Onkarnath Garg FNAE Director Indian Institute of Petroleum in Dehradun Rajkumar Ghosh Executive Director Indian Oil Andrea Gragnani Refining Product Line Director Technip Dr. Syamal Poddar President Poddar & Associates Andrea Amoroso Vice President Process Technology eni Stephany Romanow Editor Hydrocarbon Processing Michael Stockle Chief Engineer Refining Technology Foster Wheeler

Make your plans for IRPC 2012 Exhibit or Sponsor: Bill Wageneck, Vice President and Publisher, Hydrocarbon Processing at +1 (713) 520-4421 or Bill.Wageneck@GulfPub.com For more information: Teresa Wright, Director, Global Events, Gulf Publishing Company at +1 (713) 520-4475, or Teresa.Wright@GulfPub.com or go to www.HPIRPC.com

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HPIN BRIEF BILLY THINNES, TECHNICAL EDITOR

BT@HydrocarbonProcessing.com

Deloitte has released a report predicting that the future of US refining is in the hands of the independents. The report says that, by the end of 2013, over 70% of US refining capacity could be controlled by independent refiners. This is in sharp contrast to 20 years ago, when the vast majority of US refining was done by integrated oil companies.

The total value of US oil and gas mergers and acquisitions increased significantly in 2011 due to continued investment in shale plays and related infrastructure, sustained interest from foreign buyers and private equity entrants deploying capital in the energy industry, according to a new report from professional services firm PricewaterhouseCoopers. A major trend in the energy sector driving the increase in deal value throughout the year was a shift towards more investments in oil and liquid plays as natural gas prices remained depressed, hitting a 10-year low in 2011. In 2011, there were 191 deals with values greater than $50 million, accounting for $186.5 billion, a significant jump in total deal value from the $138.5 billion during 2010, which had five more announced deals. Average deal size also increased in 2011 to $977 million, a 38% jump from $706 million in 2010.

The director-general of the United Nations Industrial Development Organization recently urged countries across Asia to commit to achieving sustainable energy. “Reaching the goal of sustainable energy for all will require action by all countries and all sectors to shape the policy and investment decisions needed for a brighter energy future,” Dr. Kandeh Yumkella said. “Industrialized countries must accelerate the transition to low-emission technologies. Developing countries, many of them growing rapidly and at large scale, have the opportunity to leapfrog conventional energy options and move directly to cleaner energy alternatives that will enhance economic and social development.” The UN has set three complementary objectives to be achieved by 2030: ensure universal access to modern energy services, double energy efficiency and double the share of renewable energy in the global energy mix.

Aquaviridis has signed a commercial agreement with OriginOil to help develop the multi-phase algae production rollout at a site in Mexicali, Mexico. The two companies are heralding this as a potential model for algae sites throughout the North American Free Trade Agreement (NAFTA) region, with a focus on desert areas of the American Southwest and Mexico. OriginOil will provide its expertise to help develop growth and harvesting solutions. Aquaviridis plans to scale up from research and development to 10 acres of pilot algae production by the middle of this year. Commercial-scale production capacity is expected by the second quarter of 2013.

Brian MacDonald will soon replace Lynn Elsenhans as Sunoco CEO amid the company’s ongoing exit from the manufacturing industry and transition to a logistics and retail focus. US-based Sunoco announced last September that it plans to leave the refining business by mid-2012. Effective March 1, Mr. MacDonald will become president, CEO and director of Sunoco. He currently serves as senior vice president and chief financial officer for the company. Ms. Elsenhans will remain chairman of Sunoco and Sunoco Logistics until Sunoco’s annual shareholder’s meeting in May, at which time Mr. MacDonald will become chairman of both companies.

Eastman Chemical has agreed to acquire US-based specialty chemicals and performance materials firm Solutia, a global leader in performance materials and specialty chemicals, in a deal valued at about $4.7 billion (including debt). Eastman and Solutia share several key fundamentals, such as complementary technologies and business capabilities, a polymer science backbone, similar operating philosophies and a high performance culture, the companies said. HP

■ US demand grows for woodplastic composite US demand for wood-plastic composite and plastic lumber is projected to advance by over 13% per year to $5.4 billion in 2015, creating a market for 2.6 billion pounds of plastic. Advances will be driven by a rebound in construction expenditures from a depressed 2010 base. Growth will be further boosted by increasing consumer demand for building products made from composite and plastic lumber, instead of more traditional materials, such as natural wood. These and other trends are presented in a new study from The Freedonia Group, Inc. Decking, which was the leading application for composite and plastic lumber in 2010, will experience the most rapid demand advances through 2015. Among other applications, molding and trim, as well as windows and doors, are expected to post the quickest demand gains through 2015. Homeowners will install windows and doors made from cellular PVC and composite lumber because of their resistance to rotting and resemblance to natural wood. Demand for composite and plastic lumber in landscape and outdoor products, fencing and other applications will be promoted by increasing consumer recognition of the performance properties of these materials. Demand for wood-plastic composite lumber will post more rapid gains than that for plastic lumber through 2015, advancing over 16% annually to $2.5 billion. Gains will be driven by ongoing consumer interest in composite lumber as a substitute for natural wood products in such applications as decking and fencing. Moreover, because wood-plastic composite lumber incorporates recycled materials, it is seen as an environmentally friendly building material. Plastic lumber demand is forecast to rise nearly 11% per year to $2.8 billion in 2015. Gains will be spurred by rising consumer interest in the material because of its low maintenance properties. HP HYDROCARBON PROCESSING MARCH 2012

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HPINSIGHT Technology and market share: Challenges requiring ever-changing solutions In this issue of HPInsight, workforce, fuel quality and capacity overhang are a few of the challenges that the hydrocarbon processing industry (HPI) must address and resolve. Expansion of the refining and petrochemical industries is followed by excess supplies chasing dwindling demand. This is not a new condition; however, history and technology do modify how the HPI re-emerges from the slump.

3,000 service stations. This would be the first time that a private– sector Saudi Arabian company (Arabian Petroleum) has processed and marketed Arab crude outside the country. The 50-50 JV, Fina USA, aims to own refineries at Port Arthur (150,000 bpd), and Big Spring (60,000 bpd), both in Texas. In a similar plan, Saudi Aramco owns 50% of Star Enterprises, a 600,000-bpd refining network, and Texaco owns the other half.

Headlines from Hydrocarbon Processing, March 2002:

US refining capacity rationalization has begun. Several refiners have announced “downsizing” and/or proposed sales of assets. Chevron, Shell, Amoco and Phillips are among those refiners rationalizing. In addition, many smaller refineries have announced temporary shutdowns due to poor margins. According to a Salomon report, the US refining industry is the envy of the world in terms of light product yields. But US refiners are “tiny” by world standards. At present, 194 US refineries are operating and have an average distillation capacity under 80,000 bpd. European and Asian refineries have an average distillation capacity of 200,000 bpd. About 90 US refineries have a processing capacity under 50,000 bpd. When the rationalization round is complete, the US refining industry will emerge a smaller and more profitable industry.

What is the fuel combination to run future engines? Is zero pollution an objective that can be reached within five years, using available technologies? Yes, say representatives of automakers and engine R&D centers meeting at an international conference organized by the French Petroleum Institute (IFP). IFP claims that technologies for combustion by controlled auto-ignition in gasoline engines and homogeneous charge compression ignition for diesel engines are going to change considerably to protect the environment. IFP is developing a new approach to diesel combustion engines; it uses multiple-injection strategies. Concerns remain about fuel quality for new engines. Chemicals slump: This, too, shall pass, but when? With the significant capacity overhang prevalent today, petrochemical producers will need to closely monitor comparative international oil and domestic natural gas prices, according to a CMAI report. In 2001, weakening economies became a focal point for the global economy, accompanied by a severe downturn in petrochemical demand. Strong recovery is anticipated for 2002.

World LNG industry is growing. The world liquefied natural gas (LNG) industry is on the move, according to a new study by

More perspective on mergers and acquisitions. In the HPI, it is difficult to escape the “boom and bust” cycle. Corporate buying and selling have been extremely active for HPI companies, according to Accenture. Acquisitions can be very fruitful in bust cycles if done well. The Conoco/Phillips merger is one example of a promising and clear transaction. Edinburgh becomes the world’s first city to offer both sulfurfree unleaded gasoline and sulfur-free diesel. These fuels became available in mid-February 2002 at 18 BP service stations. The new fuels are said to be the cleanest gasoline and diesel products available in the UK. (The fuel formulations are allowed to have a maximum sulfur content of 10 ppm.) The new clean fuels, arriving six years ahead of EU legislation requirements, are produced at BP’s Grangemouth refinery in eastern Scotland.

Headlines from Hydrocarbon Processing, March 1992: The Saudis want to enter the US refining market via a joint venture (JV) involving half of Fina’s refining and marketing assets. The $1.3 billion deal would involve two Texas refineries and

The secondary fractionating tower at Sinclair Refining Corp.’s Corpus Christi, Texas refinery is eight stories tall. It weighs 90,000 lb and has 24 process trays. Petroleum Refiner 1957. HYDROCARBON PROCESSING MARCH 2012

I 11

HPINSIGHT CEDIGAZ. In 1991, LNG trade reached 78.1 billion m3 (58 metric tons), an 8% increase over 1990 levels. The Asia Pacific region was particularly active, with an 11% growth in import volume. Japan, South Korea and Taiwan increased LNG imports from Malaysian and Australian liquefaction facilities. Two new grassroots facilities are planned: The Bonny Island, Nigeria, project will supply Europe and a new LNG facility will be built in Qatar to export LNG to Japan. Eight new LNG projects will be developed by 2010 with an estimated capital cost of $30 billion. US crude oil price ‘to hit’ $20/bbl soon. US crude oil prices will rise above $20/bbl, and natural gas (NG) will increase to nearly $2/Mcf by year end. The oil and NG prices now lag behind the economy. The global oversupply of crude oil has depressed US oil and NG prices. Seasonal demand for oil and NG was reduced due to mild winter temperatures. Domestic drilling for oil and NG is the lowest in the past 50 years due to price sensitivity.

Headlines from Hydrocarbon Processing, March 1982: Natural gas pipeline to the Lower 48 from Alaska is closer than ever to a becoming reality. The Alaska natural-gas (NG) transportation system will have an initial capacity of 2 Bcfd of NG, enough to displace 400,000 bpd of crude oil for 25 to 30 years. The 745-mile Alaska segment of the project will be built and operated by a consortium of 10 US and Canadian NG companies. Oil price decontrol proves to be no ‘evil’. Decontrol of crude oil prices initiated a flurry of dire predictions. With the complete phase-out of crude oil price controls in early 1981, oil production in the Lower 48 states nullify predictions that production would decline. Without price controls, the oil industry increased drill-

ing to an all-time high. An estimated $50 billion was invested in E&P. Industry pessimists predicted that decontrol would lead to skyrocketing oil prices. In reality, the average price of a gallon of gasoline was 5¢ to 6¢ less than the peak price in 1981. With all of the improvements under decontrolling oil prices, the US still imports one third of its daily oil consumption. The nation is still very vulnerable to sudden, major disruptions of foreign oil supplies. CEFIC investigates Western Europe’s chemical industry. Western Europe’s chemical industry continues to face economic and supply/demand imbalances. The European Council of Chemical Manufacturers Federation's (CEFIC’s) view is that Western Europe is in a deep recession accompanied by structural inflation. During the 1960s, chemical production increased 15%/yr. This annual growth stabilized in the 1970s to 5%–7%. The rapid capacity expansion of the 1960s created excess production capacity. Western Europe now faces completion from Eastern countries. The Western European chemical industry must develop a new strategy based on structural changes. The new focus will be on raw material supply, processing efficiency and developing a range of products for manufacturing and distribution. The European chemical industry must find a balanced system in which profitability and productivity are depenalized. Methanol research continues to make an impact for transportation fuels. Two methanol-powered automobiles with unique prevaporized fuel system designs are undergoing a two-year test by Conoco. Company employees will test drive the methanol cars—1981 Ford Fairmonts—under normal conditions. Gasoline is used to start the engine and heat it to a set temperature before a sensor switches the engine to methanol. Because the methanol is vaporized first, there may be less cylinder wear.

Headlines from Hydrocarbon Processing, March 1972: New regulations are proposed for no-lead gasoline. The US Environmental Protection Agency (EPA) has set a July 1974 deadline by which gasoline with 91 RON or less will be lead- and phosphorous-free. Lead in regular and premium gasoline should not exceed 2 g/gal after January 1974, 1.7 g/gal after January 1975, 1.5 g/gal after January 1976, and 1.25 g/gal after January 1977. Toray has new styrene extraction process. A new process to extract styrene from cracked oil coproduced in naphtha cracking has been developed by Toray Industries. The process, STEX (styrene-extraction), is claimed to provide several advantages, such as low material cost and a simple process flow. Pyrolysis process converts waste into fuels. Occidental Petroleum has developed a new pyrolysis process that can convert municipal solid waste into low-sulfur fuels and other salable products. The new process can recover 90% of the raw materials contained in municipal trash. The process does not require hydrogenation, and it operates at atmospheric pressure. Shredded waste is mixed with pulverized coal at a 90:10 ratio. The pilot program is sponsored by the US Environmental Protection Agency. Dow Chemical Co.’ olefins production facility in Fort Saskatchewan, Alberta, Canada, expanded operating capacity to 2.4 billion lb/yr in mid-1998. Hydrocarbon Processing 1999. 12

I MARCH 2012 HydrocarbonProcessing.com

To see more headlines from 1962 to 1922, visit HydrocarbonProcessing.com.

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Repair is an Opportunity for Pump Improvement Pump Upgrades and Rerates Hydro Engineers have been improving pump performance since 1969. Many pumps operating today were designed and manufactured decades ago. Operating requirements of the plant or process may have changed and the pump is no longer operating at its BEP. By reviewing the pumpâ&#x20AC;&#x2122;s original design in relation to todayâ&#x20AC;&#x2122;s requirements, Hydroâ&#x20AC;&#x2122;s engineering staff can recommend upgrades for improving performance and extending pump life.

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

Consider procedure cards, technical books and maintenance conferences for training TABLE 1. Essential reliability library in 2012 Bloch, H. P., Improving Machinery Reliability, 3rd Ed., Gulf Publishing Company, Houston, Texas, 1998. Bloch, H. P. and F. Geitner, Machinery Failure Analysis and Troubleshooting, 3rd Ed., Gulf Publishing Company, Houston, Texas, 1998. Bloch, H. P. and F. Geitner, Machinery Component Maintenance and Repair, 3rd Ed., Gulf Publishing Company, Houston, Texas, 2004. Bloch, H. P. and F. Geitner, Major Process Equipment Maintenance and Repair, 2nd Ed., Gulf Publishing Company, Houston, Texas, 1997 Bloch, H. P. and A. Shamin, Oil Mist Handbook: Practical Applications, Fairmont Press, Lilburn, Georgia, 1998. Bloch, H. P. and C. Soares, Process Plant Machinery, 2nd Ed., Elsevier Publishing, New York and London, 1998. Bloch, H. P. and J. J. Hoefner, Reciprocating Compressors: Operation and Maintenance, Gulf Publishing Company, Houston, Texas, 1996. Bloch, H. P., Practical Guide to Compressor Technology, 2nd Ed., John Wiley & Sons, Hoboken, New Jersey, 2006; 1st Ed. available also in Spanish from McGraw-Hill. FIG. 1

Shirt-pocket checklists used at a best-of-class plant in 1989.

Because our memories are imperfect and we don’t want to risk making mistakes, we learned to jot down brief reminder notes on paper. Some people call it a cheat sheet, or a checklist. An airline pilot will find a set of checklists in the cockpit. Even after hundreds of successful takeoffs and landings, the pilot would use the checklist. Chances are that consistently using such a checklist is a condition of employment for airline cockpit crews. Errors are costly and dangerous. Due to the consequences from errors, a few best-of-class petrochemical companies issue pertinent checklists to every plant operator, as shown in Fig. 1. One best-of-class petrochemical company is Esso Chemicals—predecessor of the ExxonMobil (EM) Corp. Whenever the Esso operators are patrolling their process units, they have the checklists in their shirt pockets. Today, Esso and its successor affiliates still believe that an informed workforce is a happy, responsible and productive partner. In 2012, EM’s lube marketing branch is fulfilling a teaching and information-sharing role by organizing a lubrication-oriented maintenance symposiums that will be held in several cities in the US, Canada and Mexico.* As in previous years, the attendees will hear reliability-focused messages that are not presented elsewhere. Just as on a dozen or so occasions last year, it will be this writer’s privilege to make two different presentations at each venue. But even if you cannot attend, remember our checklist message. At best-of-class companies, unit supervisors lead by example. They, too, have these checklists with them and they will look at them every time a machine is being started or stopped or—in the case of valves—when an adjustment has to be made. Also, at best-

Bloch, H. P. and M. Singh, Practical Guide to Steam Turbine Technology, 2nd Ed., Singh New York, New York, 1995, 2009; 1st Ed. available also in Spanish from McGraw-Hill. Bloch, H. P., Practical Lubrication for Industrial Facilities, 2nd Ed., Fairmont Press, Lilburn, Georgia, 2000, 2009. Bloch, H. P. and C. Soares, Turboexpanders and Process Applications, Elsevier Publishing, New York and London, 2001. Bloch, H. P., Pump Wisdom: Problem Solving for Operators and Specialists, John Wiley & Sons, Hoboken, New Jersey, 2011. Bloch, H. P. and A. Budris, Pump User’s Handbook: Life Extension, 3rd Ed., Fairmont Press, Lilburn, Georgia, 2004, 2006, 2009. Bloch, H. P. and F. Geitner, Maximizing Machinery Uptime, Gulf Publishing Company, Houston, Texas, 2006. Bloch, H. P., Compressors and Modern Process Applications, John Wiley & Sons, Hoboken, New Jersey, 2006. Sofronas, A., Analytical Troubleshooting of Process Machinery and Pressure Vessels—Including Real-World Case Studies, John Wiley & Sons, Hoboken, New Jersey, 2006.

of-class companies, books are still within reach of operators and technical staff. The Internet is not an adequate substitute for good books. Table 1 is a partial reading list for reliability engineers. HP * Note: The Mobil Lubricants Maintenance Symposium will be held in various cities throughout North America. More information on the dates and locations for these symposiums can be found online at: www.travelhq.com/events/ exxonmobil2012/customerindex.mtc. The author is Hydrocarbon Processing’s Reliability/Equipment Editor. A practicing consulting engineer with 50 years of applicable experience, he advises process plants worldwide on failure analysis, reliability improvement and maintenance costavoidance topics. HYDROCARBON PROCESSING MARCH 2012

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expected to slow significantly after 2020 as the economy matures (Fig. 2). Although India’s population is on track to exceed China’s, its energy growth path is unlikely to replicate China’s energy intensive growth path. It will more than double its energy use to 2030, heavily based on coal, but this will still result in the consumption of some 1.3 billion toe, or just over one quarter of China’s total. There will remain a heavy reliance on higher oil exports from Middle East OPEC countries to meet demand. BP’s analysis suggests that the Middle East countries have the capability to bring on the required new

0.4

50 Shares of world primary energy, %

renewables, including biofuels, continue to be the fastest-growing sources of energy globally, rising at an annual clip of more than 8%, much quicker even than natural gas, the fastest-growing fossil fuel at about 2% per year over the period to 2030. BP chief economist Christof Rühl argues that the impact of globalization and competition will continue to deliver a remarkable convergence in energy intensity around the world (Fig. 1), a measure of energy use per unit of national economic output. The growth of unconventional supply, including US shale oil and gas, Canadian oil sands and Brazilian deepwaters, against a background of a gradual decline in oil demand, will see the Western Hemisphere become almost totally energy self-sufficient by 2030. This means that growth in the rest of the world, principally Asia, will depend increasingly on the Middle East for its growing oil requirements. Oil will continue to lose market share throughout the period, although demand for hydrocarbon liquids will still reach 103 million bpd in 2030, up by 18% from 2010. This means the world will still need to bring on enough liquids to meet the forecasted 16 million bpd of extra demand by 2030 and replace declining output from existing sources.

China. In China, growth of energy use is

0.3 China

US 0.2 World

India

0.1

0 1970

1990

2010

Oil

40 30 Coal 20 Gas 10

Nuclear

0 1970

2030

Hydro Renewables*

1990

2010

2030

Source: BP 2012 Energy Outlook 2030 * Includes biofuels

FIG. 1

Convergence of energy intensity and fuel shares.

5 Renewables* Hydro Nuclear Coal Gas Oil

4 China, billion toe

Switching. Gradual switching should see

While coal is expected to continue gaining market share in the current decade, growth will wane in the 2020–2030 decade; gas growth will remain steady and non-fossil fuels are likely to contribute nearly half of the growth after 2020. Power generation is expected to be the fastest-growing user of energy in the period to 2030, accounting for more than half the total growth in primary energy use. And it is in the power sector where the greatest changes in the fuel mix are expected. Renewables, nuclear and hydroelectric should account for more than half the growth in power generation. Energy intensity, Toe per $1,000, 2010 GDP

Global energy demand will continue to grow over the next 20 years, albeit at a slowing annual rate, fueled by economic and population growth in non-OECD countries. Increased energy efficiency and strong growth for renewable energy are also forecast in BP’s Energy Outlook 2030, which was recently released. Global energy demand is likely to grow by 39% by 2030, or 1.6% annually, almost entirely in non-OECD countries; consumption in OECD countries is expected to rise by just 4% in total over the period. Global energy will remain dominated by fossil fuels, which are forecast to account for 81% of global energy demand by 2030 down about 6% from current levels. The period should also see increased fuelswitching, with more gas and renewables use at the expense of coal and oil.

India, billion toe

BP’s energy outlook to 2030

1 0 1990

2000

2010

2020

2030

0 1990

2000

2010

2020

2030

Source: BP 2012 Energy Outlook 2030 * Includes biofuels

FIG. 2

Energy consumption growth in India and China.

HYDROCARBON PROCESSING MARCH 2012

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HPIMPACT production to meet global demand, even though the region’s energy use per capita is expected to remain more than three times as high as the rest of the non-OECD world. BP expects to see steady progress in longstanding efforts to displace oil with gas and to improve the efficiency of energy use within the region. Saudi Arabian, Iraqi and regional production of gas-related liquids will dominate supply growth as the region’s share of global oil supply rises to 34% by 2030. Transportation is likely to be the slowest-growing sector for global energy consumption; significant improvements in fuel efficiency, including hybridization of vehicles, will partly offset continued strong growth in vehicle sales in emerging markets. Hybrid vehicles (including plug-ins) offer consumer flexibility and appear capable of meeting anticipated fuel economy targets in 2030; oil is likely to account for 87% of transport sector energy use, down from 95% today, with biofuels filling most of the gap and accounting for 7% of transport sector energy use.

more aggressive policies than currently envisioned are introduced, global CO 2 emissions could begin to decline by 2030. By 2030, energy importers will need to import 40% more than they do today, but the experience will vary by region. In North America, efforts to reduce dependence on foreign supplies should show impressive results in the next couple of decades. Bolstered by supply growth from biofuels, as well as unconventional oil and gas, North America’s energy deficit will turn into a small surplus by 2030. In contrast, Europe’s energy deficit remains at current levels for oil and coal but will increase by some two-thirds for natural gas, supplied by LNG and pipelines from the former Soviet Union. China’s energy deficit across all fuels will widen by more than a factor of five and India’s (mainly oil and coal), will more than double in the period to 2030.

Emissions. Global CO2 emissions are likely to rise by about 28% by 2030. If

Filtering CO 2 from factory smokestacks is a necessary, but expensive, part

Octagonal opportunity for carbon capture

of many manufacturing processes. However, a collaborative research team from the National Institute of Standards and Technology (NIST) and the University of Delaware has gathered new insight into the performance of a material called zeolite that may stop carbon dioxide in its tracks far more efficiently than current scrubbers do. The roughly octagonal pores in zeolite SSZ-13 are like stop signs for carbon dioxide, capturing molecules of the greenhouse gas while apparently letting other substances through (Fig. 3). The material could prove to be an economical smokestack filter. Zeolites are highly porous rocks— think of a sponge made of stone—and while they occur in nature, they can be manufactured as well. Their toughness, high surface area (a gram of zeolite can have hundreds of square meters of surface in its myriad internal chambers) and ability to be reused hundreds of times make them ideal candidates for filtering gas mixtures. If an unwanted molecule in the gas mixture is found to stick to a zeolite, passing the mixture through it can scrub

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HPIMPACT the gas of many impurities, so zeolites are widely used in industrial chemistry as catalysts and filters. The team explored a zeolite created decades ago in an industrial lab and known by its technical name, SSZ-13. This zeolite, which has octagonal “windows” between its interior pore spaces, is special because it seems highly capable of filtering out CO2 from a gas mixture. “That makes SSZ-13 a promising candidate for scrubbing this greenhouse gas out of such things as factory smokestacks,” said Craig Brown, a researcher at the NIST Center for Neutron Research (NCNR). “So we explored, on an atomic level, how it does this so well.” Using neutron diffraction, the team determined that SSZ-13’s eight-sided pore windows are particularly good at attracting the long, skinny carbon dioxide molecules and holding onto their “positively charged” central carbon atoms, all the while allowing other molecules with different shapes and electronic properties to pass by unaffected. Like a stop sign, each pore halts one CO2 molecule—and each cubic centimeter of the zeolite has enough

pores to stop 0.31 grams of CO2, a quantity that the research team says makes SSZ13 highly competitive when compared to other adsorbent materials. Mr. Brown said a zeolite like SSZ-13 probably will become a prime candidate for carbon scrubbing because it also could prove more economical than other scrubbers currently used in industry. SSZ-13’s ability to attract only CO2 could mean its use would reduce the energy demands of scrubbing, which can require up to 25% of the power generated in a coal or natural gas power plant. “Many industrial zeolites attract water and carbon dioxide, which are both present in flue exhaust—meaning both molecules are, in a sense, competing for space inside the zeolite,” Mr. Brown said. “We suspect that this novel CO2 adsorption mechanism means that water is no longer competing for the same site. A zeolite that adsorbs CO2 and little else could create significant cost savings, and that’s what this one appears to do.” Mr. Brown said that his team is still collecting data to confirm this theory, and that their future efforts will con-

centrate on exploring whether SSZ-13 is equally good at separating CO2 from methane, the primary component of natural gas. CO2 is also released in significant quantities during gas extraction, and the team is hopeful that SSZ-13 can also address this problem. HP

FIG. 3

Zeolites may be a breakthrough solution for carbon capture issues.

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Process Insight:

Selecting the Best Solvent for Gas Treating

Selecting the best amine/solvent for gas treating is not a trivial task. There are a number of amines available to remove contaminants such as CO2, H2S and organic sulfur compounds from sour gas streams. The most commonly used amines are monoethanolamine (MEA), diethanolamine (DEA), and methyldiethanolamine (MDEA). Other amines include diglycolamine® (DGA), diisopropanolamine (DIPA), and triethanolamine (TEA). Mixtures of amines can also be used to customize or optimize the acid gas recovery. Temperature, pressure, sour gas composition, and purity requirements for the treated gas must all be considered when choosing the most appropriate amine for a given application.

Primary Amines

Mixed Solvents

dŚĞ ƉƌŝŵĂƌǇ ĂŵŝŶĞ D ƌĞŵŽǀĞƐ ďŽƚŚ KϮ ĂŶĚ ,Ϯ^ ĨƌŽŵ ƐŽƵƌ ŐĂƐ ĂŶĚ ŝƐ ĞīĞĐƟǀĞ Ăƚ ůŽǁ ƉƌĞƐƐƵƌĞ͘ ĞƉĞŶĚŝŶŐ ŽŶ ƚŚĞ ĐŽŶĚŝƟŽŶƐ͕ D ĐĂŶ ƌĞŵŽǀĞ ,Ϯ^ ƚŽ ůĞƐƐ ƚŚĂŶ ϰ ƉƉŵǀ ǁŚŝůĞ ƌĞŵŽǀŝŶŐ KϮ ƚŽ ůĞƐƐ ƚŚĂŶ ϭϬϬ ƉƉŵǀ͘ D ƐǇƐƚĞŵƐ ŐĞŶĞƌĂůůǇ ƌĞƋƵŝƌĞ Ă ƌĞĐůĂŝŵĞƌ ƚŽ ƌĞŵŽǀĞ ĚĞŐƌĂĚĞĚ ƉƌŽĚƵĐƚƐ ĨƌŽŵ ĐŝƌĐƵůĂƟŽŶ͘ dǇƉŝĐĂů ƐŽůƵƟŽŶ ƐƚƌĞŶŐƚŚ ƌĂŶŐĞƐ ĨƌŽŵ ϭϬ ƚŽ ϮϬ ǁĞŝŐŚƚ й ǁŝƚŚ Ă ŵĂǆŝŵƵŵ ƌŝĐŚ ůŽĂĚŝŶŐ ŽĨ Ϭ͘ϯϱ ŵŽůĞ ĂĐŝĚ ŐĂƐͬŵŽůĞ D ͘ ' Π ŝƐ ĂŶŽƚŚĞƌ ƉƌŝŵĂƌǇ ĂŵŝŶĞ ƚŚĂƚ ƌĞŵŽǀĞƐ KϮ͕ ,Ϯ^͕ K^͕ ĂŶĚ ŵĞƌĐĂƉƚĂŶƐ͘ dǇƉŝĐĂů ƐŽůƵƟŽŶ ƐƚƌĞŶŐƚŚƐ ĂƌĞ ϱϬͲϲϬ ǁĞŝŐŚƚ й͕ ǁŚŝĐŚ ƌĞƐƵůƚ ŝŶ ůŽǁĞƌ ĐŝƌĐƵůĂƟŽŶ ƌĂƚĞƐ ĂŶĚ ůĞƐƐ ĞŶĞƌŐǇ ƌĞƋƵŝƌĞĚ ĨŽƌ ƐƚƌŝƉƉŝŶŐ ĂƐ ĐŽŵƉĂƌĞĚ ǁŝƚŚ D ͘ ' ĂůƐŽ ƌĞƋƵŝƌĞƐ ƌĞĐůĂŝŵŝŶŐ ƚŽ ƌĞŵŽǀĞ ƚŚĞ ĚĞŐƌĂĚĂƟŽŶ ƉƌŽĚƵĐƚƐ͘

Secondary Amines

/Ŷ ĐĞƌƚĂŝŶ ƐŝƚƵĂƟŽŶƐ͕ ƚŚĞ ƐŽůǀĞŶƚ ĐĂŶ ďĞ ͞ĐƵƐƚŽŵŝǌĞĚ͟ ƚŽ ŽƉƟŵŝǌĞ ƚŚĞ ƐǁĞĞƚĞŶŝŶŐ ƉƌŽĐĞƐƐ͘ &Žƌ ĞǆĂŵƉůĞ͕ ĂĚĚŝŶŐ Ă ƉƌŝŵĂƌǇ Žƌ ƐĞĐŽŶĚĂƌǇ ĂŵŝŶĞ ƚŽ D ĐĂŶ ŝŶĐƌĞĂƐĞ ƚŚĞ ƌĂƚĞ ŽĨ KϮ ĂďƐŽƌƉƟŽŶ ǁŝƚŚŽƵƚ ĐŽŵƉƌŽŵŝƐŝŶŐ ƚŚĞ ĂĚǀĂŶƚĂŐĞƐ ŽĨ D ͘ DŽƌĞ ĐŽŵŵŽŶ ŝŶ ƚŽĚĂǇ͛Ɛ ŵĂƌŬĞƚ ŝƐ ƚŚĞ ĂĚĚŝƟŽŶ ŽĨ ƉŝƉĞƌĂǌŝŶĞ ƚŽ D ƐŽůƵƟŽŶƐ ĨŽƌ KϮ ƌĞŵŽǀĂů Žƌ ƉŽƐƐŝďůǇ ĂŶ ĂĐŝĚ ƚŽ ĂŝĚ ŝŶ ƌĞŐĞŶĞƌĂƚŽƌ ƉĞƌĨŽƌŵĂŶĐĞ ĨŽƌ ƚŚĞ ůĞĂŶ ƐŽůǀĞŶƚ͘ DĂŶǇ ƉůĂŶƚƐ ƵƟůŝǌĞ Ă ŵŝǆƚƵƌĞ ŽĨ ĂŵŝŶĞ ǁŝƚŚ ƉŚǇƐŝĐĂů ƐŽůǀĞŶƚƐ͘ ^h>&/EK> ŝƐ Ă ůŝĐĞŶƐĞĚ ƉƌŽĚƵĐƚ ĨƌŽŵ ^ŚĞůů Kŝů WƌŽĚƵĐƚƐ ƚŚĂƚ ĐŽŵďŝŶĞƐ ĂŶ ĂŵŝŶĞ ǁŝƚŚ Ă ƉŚǇƐŝĐĂů ƐŽůǀĞŶƚ͘ ĚǀĂŶƚĂŐĞƐ ŽĨ ƚŚŝƐ ƐŽůǀĞŶƚ ĂƌĞ ŝŶĐƌĞĂƐĞĚ ŵĞƌĐĂƉƚĂŶ ƉŝĐŬƵƉ͕ ůŽǁĞƌ ƌĞŐĞŶĞƌĂƟŽŶ ĞŶĞƌŐǇ͕ ĂŶĚ ƐĞůĞĐƟǀŝƚǇ ƚŽ ,Ϯ^͘

Choosing the Best Alternative

dŚĞ ƐĞĐŽŶĚĂƌǇ ĂŵŝŶĞ ƌĞŵŽǀĞƐ ďŽƚŚ KϮ ĂŶĚ ,Ϯ^ ďƵƚ ŐĞŶĞƌĂůůǇ ƌĞƋƵŝƌĞƐ ŚŝŐŚĞƌ ƉƌĞƐƐƵƌĞ ƚŚĂŶ D ƚŽ ŵĞĞƚ ŽǀĞƌŚĞĂĚ ƐƉĞĐŝĮĐĂƟŽŶƐ͘ ĞĐĂƵƐĞ ŝƐ Ă ǁĞĂŬĞƌ ĂŵŝŶĞ ƚŚĂŶ D ͕ ŝƚ ƌĞƋƵŝƌĞƐ ůĞƐƐ ĞŶĞƌŐǇ ĨŽƌ ƐƚƌŝƉƉŝŶŐ͘ dǇƉŝĐĂů ƐŽůƵƟŽŶ ƐƚƌĞŶŐƚŚ ƌĂŶŐĞƐ ĨƌŽŵ Ϯϱ ƚŽ ϯϱ ǁĞŝŐŚƚ й ǁŝƚŚ Ă ŵĂǆŝŵƵŵ ƌŝĐŚ ůŽĂĚŝŶŐ ŽĨ Ϭ͘ϯϱ ŵŽůĞͬŵŽůĞ͘ /W ŝƐ Ă ƐĞĐŽŶĚĂƌǇ ĂŵŝŶĞ ƚŚĂƚ ĞǆŚŝďŝƚƐ ƐŽŵĞ ƐĞůĞĐƟǀŝƚǇ ĨŽƌ ,Ϯ^ ĂůƚŚŽƵŐŚ ŝƚ ŝƐ ŶŽƚ ĂƐ ƉƌŽŶŽƵŶĐĞĚ ĂƐ ĨŽƌ ƚĞƌƟĂƌǇ ĂŵŝŶĞƐ͘ /W ĂůƐŽ ƌĞŵŽǀĞƐ K^͘ ^ŽůƵƟŽŶƐ ĂƌĞ ůŽǁ ŝŶ ĐŽƌƌŽƐŝŽŶ ĂŶĚ ƌĞƋƵŝƌĞ ƌĞůĂƟǀĞůǇ ůŽǁ ĞŶĞƌŐǇ ĨŽƌ ƌĞŐĞŶĞƌĂƟŽŶ͘ dŚĞ ŵŽƐƚ ĐŽŵŵŽŶ ĂƉƉůŝĐĂƟŽŶƐ ĨŽƌ /W ĂƌĞ ŝŶ ƚŚĞ /WΠ ĂŶĚ ^h>&/EK>Π ƉƌŽĐĞƐƐĞƐ͘

Tertiary Amines ƚĞƌƟĂƌǇ ĂŵŝŶĞ ƐƵĐŚ ĂƐ D ŝƐ ŽŌĞŶ ƵƐĞĚ ƚŽ ƐĞůĞĐƟǀĞůǇ ƌĞŵŽǀĞ ,Ϯ^͕ ĞƐƉĞĐŝĂůůǇ ĨŽƌ ĐĂƐĞƐ ǁŝƚŚ Ă ŚŝŐŚ KϮ ƚŽ ,Ϯ^ ƌĂƟŽ ŝŶ ƚŚĞ ƐŽƵƌ ŐĂƐ͘ KŶĞ ďĞŶĞĮƚ ŽĨ ƐĞůĞĐƟǀĞ ĂďƐŽƌƉƟŽŶ ŽĨ ,Ϯ^ ŝƐ Ă ůĂƵƐ ĨĞĞĚ ƌŝĐŚ ŝŶ ,Ϯ^͘ D ĐĂŶ ƌĞŵŽǀĞ ,Ϯ^ ƚŽ ϰ ƉƉŵ ǁŚŝůĞ ŵĂŝŶƚĂŝŶŝŶŐ Ϯй Žƌ ůĞƐƐ KϮ ŝŶ ƚŚĞ ƚƌĞĂƚĞĚ ŐĂƐ͕ ƚŚƵƐ ƵƐŝŶŐ ƌĞůĂƟǀĞůǇ ůĞƐƐ ĞŶĞƌŐǇ ĨŽƌ ƌĞŐĞŶĞƌĂƟŽŶ ƚŚĂŶ ƚŚĂƚ ĨŽƌ ͘ ,ŝŐŚĞƌ ǁĞŝŐŚƚ ƉĞƌĐĞŶƚ ĂŵŝŶĞ ĂŶĚ ůĞƐƐ KϮ ĂďƐŽƌďĞĚ ƌĞƐƵůƚƐ ŝŶ ůŽǁĞƌ ĐŝƌĐƵůĂƟŽŶ ƌĂƚĞƐ ĂƐ ǁĞůů͘ dǇƉŝĐĂů ƐŽůƵƟŽŶ ƐƚƌĞŶŐƚŚƐ ĂƌĞ ϰϬͲϱϬ ǁĞŝŐŚƚ й ǁŝƚŚ Ă ŵĂǆŝŵƵŵ ƌŝĐŚ ůŽĂĚŝŶŐ ŽĨ Ϭ͘ϱϱ ŵŽůĞͬŵŽůĞ͘ ĞĐĂƵƐĞ D ŝƐ ŶŽƚ ƉƌŽŶĞ ƚŽ ĚĞŐƌĂĚĂƟŽŶ͕ ĐŽƌƌŽƐŝŽŶ ŝƐ ůŽǁ ĂŶĚ Ă ƌĞĐůĂŝŵĞƌ ŝƐ ƵŶŶĞĐĞƐƐĂƌǇ͘ KƉĞƌĂƟŶŐ ƉƌĞƐƐƵƌĞ ĐĂŶ ƌĂŶŐĞ ĨƌŽŵ ĂƚŵŽƐƉŚĞƌŝĐ͕ ƚǇƉŝĐĂů ŽĨ ƚĂŝů ŐĂƐ ƚƌĞĂƟŶŐ ƵŶŝƚƐ͕ ƚŽ ŽǀĞƌ ϭ͕ϬϬϬ ƉƐŝĂ͘

'ŝǀĞŶ ƚŚĞ ǁŝĚĞ ǀĂƌŝĞƚǇ ŽĨ ŐĂƐ ƚƌĞĂƟŶŐ ŽƉƟŽŶƐ͕ Ă ƉƌŽĐĞƐƐ ƐŝŵƵůĂƚŽƌ ƚŚĂƚ ĐĂŶ ĂĐĐƵƌĂƚĞůǇ ƉƌĞĚŝĐƚ ƐǁĞĞƚĞŶŝŶŐ ƌĞƐƵůƚƐ ŝƐ Ă ŶĞĐĞƐƐŝƚǇ ǁŚĞŶ ĂƩĞŵƉƟŶŐ ƚŽ ĚĞƚĞƌŵŝŶĞ ƚŚĞ ďĞƐƚ ŽƉƟŽŶ͘ WƌŽDĂǆΠ ŚĂƐ ďĞĞŶ ƉƌŽǀĞŶ ƚŽ ĂĐĐƵƌĂƚĞůǇ ƉƌĞĚŝĐƚ ƌĞƐƵůƚƐ ĨŽƌ ŶƵŵĞƌŽƵƐ ƉƌŽĐĞƐƐ ƐĐŚĞŵĞƐ͘ ĚĚŝƟŽŶĂůůǇ͕ WƌŽDĂǆ ĐĂŶ ƵƟůŝǌĞ Ă ƐĐĞŶĂƌŝŽ ƚŽŽů ƚŽ ƉĞƌĨŽƌŵ ĨĞĂƐŝďŝůŝƚǇ ƐƚƵĚŝĞƐ͘ dŚĞ ƐĐĞŶĂƌŝŽ ƚŽŽů ŵĂǇ ďĞ ƵƐĞĚ ƚŽ ƐǇƐƚĞŵĂƟĐĂůůǇ ǀĂƌǇ ƐĞůĞĐƚĞĚ ƉĂƌĂŵĞƚĞƌƐ ŝŶ ĂŶ ĞīŽƌƚ ƚŽ ĚĞƚĞƌŵŝŶĞ ƚŚĞ ŽƉƟŵƵŵ ŽƉĞƌĂƟŶŐ ĐŽŶĚŝƟŽŶƐ ĂŶĚ ƚŚĞ ĂƉƉƌŽƉƌŝĂƚĞ ƐŽůǀĞŶƚ͘ dŚĞƐĞ ƐƚƵĚŝĞƐ ĐĂŶ ĚĞƚĞƌŵŝŶĞ ƌŝĐŚ ůŽĂĚŝŶŐ͕ ƌĞďŽŝůĞƌ ĚƵƚǇ͕ ĂĐŝĚ ŐĂƐ ĐŽŶƚĞŶƚ ŽĨ ƚŚĞ ƐǁĞĞƚ ŐĂƐ͕ ĂŵŝŶĞ ůŽƐƐĞƐ͕ ƌĞƋƵŝƌĞĚ ĐŝƌĐƵůĂƟŽŶ ƌĂƚĞ͕ ƚǇƉĞ ŽĨ ĂŵŝŶĞ Žƌ ƉŚǇƐŝĐĂů ƐŽůǀĞŶƚ͕ ǁĞŝŐŚƚ ƉĞƌĐĞŶƚ ŽĨ ĂŵŝŶĞ͕ ĂŶĚ ŽƚŚĞƌ ƉĂƌĂŵĞƚĞƌƐ͘ WƌŽDĂǆ ĐĂŶ ŵŽĚĞů ǀŝƌƚƵĂůůǇ ĂŶǇ ŇŽǁ ƉƌŽĐĞƐƐ Žƌ ĐŽŶĮŐƵƌĂƟŽŶ ŝŶĐůƵĚŝŶŐ ŵƵůƟƉůĞ ĐŽůƵŵŶƐ͕ ůŝƋƵŝĚ ŚǇĚƌŽĐĂƌďŽŶ ƚƌĞĂƟŶŐ͕ ĂŶĚ ƐƉůŝƚ ŇŽǁ ƉƌŽĐĞƐƐĞƐ͘ /Ŷ ĂĚĚŝƟŽŶ͕ WƌŽDĂǆ ĐĂŶ ĂĐĐƵƌĂƚĞůǇ ŵŽĚĞů ĐĂƵƐƟĐ ƚƌĞĂƟŶŐ ĂƉƉůŝĐĂƟŽŶƐ ĂƐ ǁĞůů ĂƐ ƉŚǇƐŝĐĂů ƐŽůǀĞŶƚ ƐǁĞĞƚĞŶŝŶŐ ǁŝƚŚ ƐŽůǀĞŶƚƐ ƐƵĐŚ ĂƐ ŽĂƐƚĂů 'ZΠ͕ ŵĞƚŚĂŶŽů͕ ĂŶĚ EDW͘ &Žƌ ŵŽƌĞ ŝŶĨŽƌŵĂƟŽŶ ĂďŽƵƚ WƌŽDĂǆ ĂŶĚ ŝƚƐ ĂďŝůŝƚǇ ƚŽ ĚĞƚĞƌŵŝŶĞ ƚŚĞ ĂƉƉƌŽƉƌŝĂƚĞ ƐŽůǀĞŶƚ ĨŽƌ Ă ŐŝǀĞŶ ƐĞƚ ŽĨ ĐŽŶĚŝƟŽŶƐ͕ ĐŽŶƚĂĐƚ ƌǇĂŶ ZĞƐĞĂƌĐŚ Θ ŶŐŝŶĞĞƌŝŶŐ͘

WƌŽDĂǆΠ ƉƌŽĐĞƐƐ ƐŝŵƵůĂƟŽŶ ƐŽŌǁĂƌĞ ďǇ ƌǇĂŶ ZĞƐĞĂƌĐŚ Θ ŶŐŝŶĞĞƌŝŶŐ͕ /ŶĐ͘ ŶŐŝŶĞĞƌŝŶŐ ^ŽůƵƟŽŶƐ ĨŽƌ ƚŚĞ Kŝů͕ 'ĂƐ͕ ZĞĮŶŝŶŐ Θ ŚĞŵŝĐĂů /ŶĚƵƐƚƌŝĞƐ͘ ƐĂůĞƐΛďƌĞ͘ĐŽŵ ǁǁǁ͘ďƌĞ͘ĐŽŵ ϵϳϵ ϳϳϲͲϱϮϮϬ h^ ϴϬϬ ϳϳϲͲϱϮϮϬ Select 71 at www.HydrocarbonProcessing.com/RS

HPIN CONSTRUCTION HELEN MECHE, ASSOCIATE EDITOR HM@HydrocarbonProcessing.com

North America UOP LLC, a Honeywell company, will provide technology to produce propylene at Dow Chemical Co.’s production site in Texas. Dow Texas Operations will use Honeywell UOP C3 Oleflex technology in a new propane dehydrogenation unit to convert shale-gas-derived propane to propylene. The facility will produce 750,000 metric tpy of polymer-grade propylene. The unit is scheduled to start up in 2015. It will reportedly be the first of its kind in the US and the largest single-train propane dehydrogenation plant in North America. Since the technology was commercialized in 1990, Honeywell’s UOP has commissioned nine C3 Oleflex units for on-purpose propylene production, with the 10th unit scheduled to start up in Russia in 2012. KBR will provide construction services for ExxonMobil’s new synthetics lubricant base-stock facility to be built at ExxonMobil’s refinery and chemical plant complex in Baytown, Texas. When completed in 2013, the facility will produce ExxonMobil Chemical’s high-viscosity SpectraSyn Elite metallocene polyalphaolefin (PAO) base stock. KBR’s scope of work for the Baytown plant includes site work, civil, structural, pipe, electrical, instrumentation and mechanical installation, as well as test and checkout services. The award by Technip USA, the prime contractor responsible for this new facility’s engineering, procurement and construction (EPC), follows the construction of ExxonMobil’s Flare Gas Recovery Project in Beaumont, Texas, for which KBR and Technip successfully collaborated to execute a complete EPC package. HollyFrontier Corp. (HFC) plans to expand capacity at its 31,000-bpd Woods Cross, Utah, refinery to 45,000 bpd with an expected completion in late 2014. The expansion includes the relocation/ revamp of crude, fluid catalytic-cracking and polymerization units from a subsidiary of Western Refining Inc.’s Bloomfield, New Mexico, refinery to Woods Cross. It also comprises an expansion of the Woods Cross diesel hydrotreater and investment in associated utilities and offsites. HFC has an

agreement with Western Refining to purchase the Bloomfield units. HFC expects incremental yields from the expansion project to be approximately 60% gasoline and 40% diesel. The expansion cost is estimated to be approximately $225 million, with an expected payback period of less than two years. In conjunction with the expansion, HollyFrontier signed a 10-year, 20,000-bpd crude-oil supply agreement with Newfield Exploration Co. This agreement, which begins once the expansion is complete, will supply black- and yellow-wax crude oil produced in the nearby Uinta Basin region to the Woods Cross refinery, which has capacity to process approximately 10,000 bpd of these crudes. When the expansion is complete, the Woods Cross refinery will be able to process approximately 24,000 bpd of waxy Utah crudes. This expansion, crude-oil supply agreement and expected completion timeline are subject to HollyFrontier successfully obtaining the necessary permits and regulatory approvals. ONEOK Partners, L.P.’s new 100 million-cfd natural gas processing Garden Creek plant in eastern McKenzie County, North Dakota, is now operational and serving producers in the Bakken Shale region. The company plans to invest approximately $1.5 billion to $1.8 billion for growth projects in the Bakken Shale between now and 2014 in its natural gas gathering and processing, and natural gas liquids (NGL) businesses. In addition to the Garden Creek plant, these investments include the construction of the Bakken pipeline, an approximately 500-mile NGL pipeline and two additional 100 millioncfd natural gas processing facilities—the Stateline I and Stateline II plants in western Williams County, North Dakota. The Bakken pipeline is expected to be completed by the first half of 2013. The Stateline I and Stateline II plants are anticipated to be completed by the third quarter of 2012 and the first half of 2013, respectively. CB&I has been awarded a contract, valued in excess of $750 million, by Imperial Oil Resources Ventures Ltd. for work on the Kearl Expansion Project in Alberta, Canada.

CB&I’s work scope on the expansion project includes the engineering, procurement, module assembly and construction of a second bitumen-extraction plant, along with froth tank farms, multiple storage tanks and six froth-settling units. Methanex Corp. is planning to move one of its idle methanol plants in Chile to some land that it has secured in Geismar, Louisiana. Site-specific engineering has begun and the plant is expected to be operational in the second half of 2014. Renewable Manufacturing Gateway (RMG) and Aither Chemicals LLC have agreed to collaborate to finance and build a large chemical plant using Aither’s ethane catalytic-cracker technology. With an investment of $750 million over the next five years, the project is expected to create over 2,000 construction jobs, 200 permanent direct production jobs, and many thousand indirect jobs in the Tri-State Region (western Pennsylvania, eastern Ohio and northern West Virginia). It is anticipated to generate $463 million in annual sales by 2016. Today the production of ethane-derived petrochemicals utilizes steam-cracking technology developed in West Virginia by Union Carbide Corp. in the 1920s. The birth of this technology reportedly helped

Trend analysis forecasting Hydrocarbon Processing maintains an extensive database of historical HPI project information. The Boxscore Database is a 35-year compilation of projects by type, operating company, licensor, engineering/constructor, location, etc. Many companies use the historical data for trending or sales forecasting. The historical information is available in comma-delimited or Excel® and can be custom sorted to suit your needs. The cost depends on the size and complexity of the sort requested. 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 needed and receive a prompt cost quotation. Contact: Lee Nichols P.O. Box 2608, Houston, Texas 77252-2608 713-525-4626 • Lee.Nichols@GulfPub.com HYDROCARBON PROCESSING MARCH 2012

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HPIN CONSTRUCTION Union Carbide become one of the world’s largest producers of ethane-derived chemicals, and a leading manufacturer of polyethylene. Aither, also a West Virginia company, was formed in 2010 by accomplished former Union Carbide and Dow technologists and business leaders. Aither’s mission is to convert ethane to high-value chemicals. Its technology uses a patent-pending catalytic-cracking method instead of steam cracking to make ethylene.

The company will then convert the ethylene to higher-value chemicals that are easier to ship to customers locally and worldwide.

South America Exterran Holdings, Inc., has a contract from Petrobras for the design, fabrication and sale of a major natural gas processing and treating facility in Itau, Bolivia. The facility is a dew-point gas processing plant that can process 200 million scfd of natural

CREATING VALUE IN EVERY PHASE WorleyParsons provides a comprehensive range of reﬁnery and petrochemicals services through all phases of the asset lifecycle, and has been doing so for over 60 years. reﬁning@worleyparsons.com petrochemicals@worleyparsons.com

With over 2,100 cumulative reﬁning and petrochemicals projects, our customers can be assured that critical issues such as sustainability; health, safety and environment; budget; schedule; quality; operating reliability; and technical integrity are optimized in all phases, from Identify through to Operate. Visit with us after the NPRA (AFPM) Annual Meeting in San Diego, California to learn more. March 11, 2012 | Manchester Grand Hyatt | Suite 3833 | 8:00 PM | One Night Only

Select 154 at www.HydrocarbonProcessing.com/RS 22

gas. Exterran is the project’s primary contractor, providing equipment, engineering, procurement and construction products, along with services and project management. The plant is expected to start up during 2013.

Europe BASF will build a single-train 300,000 metric-tpy production plant for toluene diisocyanate (TDI) at its site in Ludwigshafen, Germany. It will also expand additional plants for TDI precursors by constructing a new hydrogen-chloride recycling plant, as well as expanding plants for nitric acid, chlorine and synthesis gas at the Ludwigshafen site. The company also plans to expand the site’s aromatics complex for the supply of toluene. Total investment, including the Ludwigshafen site’s required infrastructure, will be approximately €1 billion. Approximately 200 additional jobs will be created. Production will start at the end of 2014. When this new plant goes onstream, BASF plans to close down its 80,000 metric-tpy TDI production plant in Schwarzheide, Germany. Technip was awarded a lump-sum turnkey contract, worth more than €900 million (Technip share around €600 million), by Lukoil Neftochim Burgas AD, a subsidiary of OAO LUKOIL. The contract includes Phase 1 engineering, procurement and construction (EPC) of a heavy-residue hydrocracking complex for the Lukoil Neftochim Burgas refinery in Burgas, Bulgaria. This contract covers the detail engineering, procurement of equipment and material, construction, pre-commissioning and commissioning of a 2.5 million-tpy vacuum-residue hydrocracker based on the Axens H-Oil process, as well as an amine-regeneration unit, a sour-water stripper, hydrogen production units, and utilities and offsite upgrading. The contract, which is scheduled to be completed by the end of January 2015, follows the successful execution of the front-end engineering design (FEED) completed by Technip in the first quarter of 2010, and the detailed engineering and procurement-services contract won at the beginning of 2011. A subsidiary of Foster Wheeler AG’s Global Engineering and Construction Group has been awarded a feasibility study by Albanian Refining & Marketing of Oil Sh.a. (ARMO) relating to the modernization of two refineries, located at Ballsh and Fier in Albania.

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U Creating Effective Business Strategy/Decisions U Increasing Return on Investments U Enhancing Returns on Acquisitions/Divestitures U Reducing Strategic/Capital/Market/Investment Risk U Enhancing Yields U Creating Effective Response to Crude/ Feedstock/Product Markets U Improving Financial Performance

U Increasing Return on Capital Investment U Rationalising/Optimising Environmental Compliance Capital Expenditures U Reducing Capital Risk

Operating Challenges U Improving Yield U Increasing Availability U Reducing Maintenance Costs U Improving Safety Performance U Implementing/Improving U Managing Operational Risk Behaviour-based Reliability U Improving Supply Chain Performance

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HPIN CONSTRUCTION ARMO intends to modernize the existing refineries to restore production to the original design capacity and produce transportation fuels in line with current European Union regulations. The study is expected to be completed by mid-2012.

Middle East Qatargas has signed an engineering, procurement and construction (EPC) contract with Samsung Engineering Ltd. for a diesel hydrotreater (DHT) unit that will treat 54,000 bpsd of diesel, from highsulfur into ultra-low-sulfur diesel fuel, at the Laffan refinery. At a cost of around QR 350 million, the DHT unit, which is aiming to produce diesel with less than 10 ppm of sulfur content with the Euro 5 specification, will be built and integrated into the existing Laffan refinery by 2014. The hydrotreater will process straight-run light gasoil (LGO) feedstock from the existing Laffan refinery 1 (LR1) and the second planned refinery (LR2). The DHT unit has a processing capacity of 54,000 bpsd. Until the second refinery is operational, the unit will run at 50% of its designed capacity. Occasionally, when

the existing kerosine hydrotreater is shut down either for maintenance or catalyst replacement, the DHT unit will also be able to treat the straight-run LGO from the kerosine condensate fractionation unit. The DHT unit will be installed inside the plot of LR1, which is located in Ras Laffan Industrial City in Qatar. The DHT unit project is being developed by Laffan Refinery Co., which is operated by Qatargas. The shareholders in the joint venture include Qatar Petroleum (84%), Total (10%), Idemitsu (2%), Cosmo (2%), Mitsui (1%) and Marubeni (1%). Saudi Aramco and Sinopec have agreed to form a joint venture related to the ongoing development of Yanbu Aramco Sinopec Refining Co., Ltd. (YASREF), formerly the Red Sea Refining Co. The agreement brings together two world-class companies to complete the construction and operate a highly competitive full-conversion refinery in Yanbu. The YASREF project involves construction of a new grassroots refinery on the Yanbu site covering over 5.2 million

Asset Longevity Plant & Pipeline Performance

square meters. With construction well underway, the project is on schedule with 10% of construction completed. The YASREF refinery is scheduled to be operational in the second half of 2014. The refinery will process 400,000 bpd of Arabian Heavy crude oil and produce high-quality transportation fuels. It is set to produce 90,000 bpd of gasoline, 263,000 bpd of ultra-low-sulfur diesel, along with byproducts consisting of 6,200 metric tpd of petroleum coke and 1,200 metric tpd of sulfur. Saudi Aramco’s extensive and integrated hydrocarbon facilities in Yanbu will be used to supply crude-oil feedstock to YASREF and to export transportation fuels. The refinery project includes megaprocessing units, utilities and interconnecting piping, associated crude-oil and refined product storage, as well as the offsite facilities necessary to support safe and efficient refinery operation. It is expected that approximately 60% of the total project value will be spent in Saudi Arabia through detailed engineering executed in local design offices, material procurement from local manufacturers and suppliers, and the utilization of Saudi construction companies. In addition, new technology consisting of petroleum coking will be implemented as part of this project.

Asia Pacific

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Select 155 at www.HydrocarbonProcessing.com/RS 24

Total is consolidating its positions in petrochemicals in Asia with a new expansion and upgrading project for the Daesan complex in South Korea, which the group owns with Samsung as part of the Samsung Total Petrochemicals 50/50 joint venture. With costs approaching $1.8 billion, the project calls for the construction of a second aromatics unit and an ethylene-vinyl acetate (EVA) copolymer unit. The new aromatics unit will produce around 1 million metric tpy of paraxylene and 420,000 metric tpy of benzene. The new EVA unit will produce 240,000 metric tpy of ethylene-vinyl acetate copolymers. With the completion of the aromatics unit in September 2014 and the upgrade of existing paraxylene capacity in 2012, total paraxylene production capacity will be increased to 1.76 million metric tons. Technip has a services contract with BP Zhuhai Chemical Co. Ltd., a joint venture of BP and Zhuhai Port Co. Ltd., for the front-end design of a new world-scale purified terephthalic acid (PTA) plant at their

HPIN CONSTRUCTION Zhuhai site in Guangdong Province, China. This contract will be developed by an integrated Technip/client team. The new plant (Zhuhai 3) will have a capacity of 1,250,000 tpy. It will use BP’s latest proprietary technology and is expected to come onstream during 2014, subject to final approvals from shareholders and related government authorities. Technip’s operating center in Rome, Italy, will execute the contract, which is scheduled to be completed in the first half of 2012. KBR has a contract to perform a dynamic simulation study for the complete steam system at Matix Fertilizers and Chemicals, Ltd.’s grassroots Panagarh Fertilizer Complex in West Bengal, India. KBR will also deliver an operator training simulator (OTS) for the ammonia plant system to validate the plant’s controls and safety logic design, and provide initial and ongoing training to operators. For the dynamic simulation study, KBR will develop a detailed model of the complex’s steam system, perform agreed cases and scenarios, and provide a final study

report inclusive of all results and recommendations. This simulation will enable Matix to validate the design of critical plant systems, including auxiliary boilers, letdown valves and controls. It will also validate operating procedures, such as startups, turndowns and the handling of process upsets. KBR will also deliver an OTS system for one of the largest-capacity ammonia units at Matix to ensure a safe, fast and efficient startup. The OTS system will ensure best energy consumption in the ammonia process unit and support continued profitable and sustained ammonia plant operations. INPEX CORP. and Total S.A. (Ichthys owners) have notified the joint venture formed by KBR, JGC Corp. and Chiyoda Corp. (JKC JV) that the final investment decision for the Ichthys LNG project has been achieved. As part of this notification, the Ichthys owners have issued a notification of award to the JKC JV for the Ichthys LNG project’s engineering, procurement and construction (EPC). The Ichthys owners and the JKC JV have executed a formal letter of award for the project’s initial EPC activities. The com-

Hindustan Petroleum Corporation Limited

pletion and execution of the formal EPC contract is expected in the coming weeks. A subsidiary of Foster Wheeler’s Global Engineering and Construction Group has a contract with Petron Corp. for the Petron Refinery Master Plan-2 project in Bataan, Philippines. Foster Wheeler will execute detailed engineering and procurement services for the delayed coker unit (DCU), including the engineering and material supply of two double-fired terrace wall coker heaters. The DCU will have a design capacity of 37,500 bpsd. It is a key part of this significant refinery upgrade. This award follows an earlier award for the process design package and technology license for the DCU, which will use the company’s leading Selective Yield Delayed Coking (SYDEC) process. Foster Wheeler’s SYDEC process is a thermal conversion process used by refiners worldwide to upgrade heavy residue feed and process it into high-value transport fuels. The SYDEC process can be designed to maximize clean liquid yields while minimizing fuel coke yields from high-sulfur residues.

You can make a mark in the world of oil

(A Government of India Enterprise) Regd. Office: Petroleum House, 17, Jamshedji Tata Road, Mumbai-400 020.

Requires Experienced Professionals for Refinery Operations Without refining, the rich resources of crude petroleum of nature would remain latent. Value-added products from crude petroleum like petrol, diesel, kerosene, liquefied petroleum gas, naphtha and many more products would not be available for growth and development of a nation. An Opportunity beckons Professionals having experience in Petroleum Refineries / Petrochemical / Fertilizer Industry, for bright career prospects at HPCL Government of India Enterprise with a Navratna Status, and a Fortune 500 and Forbes 2000 company, with an annual turnover of ` 1,32,670 crores. HPCL refineries upgrade the crude petroleum into many value-added products and over 300 grades of lubricants, specialties and greases. The Lubricating Oils Refinery set up at Mumbai is largest lube refinery in India producing superior quality lube base oils. The production facilities at both Mumbai and Vishakhapatnam refineries are upgraded to produce green fuels like unleaded Euro III & Euro IV Petrol and ultra-low sulphur Euro III Diesel. Project for production of Euro IV Diesel is also expected to commence shortly. Our refineries have been benchmarked by an international agency for various performance parameters. Towards, achieving HPCL's vision of becoming World Class Energy Company, we are looking for competent & experienced professionals with excellent track record for our refinery operations.

Position

Discipline

No of Posts

Rotary Engg / Mechanical Project Planning Engg/ Mechanical Design Engg / Mechanical Construction Engg

Mechanical

Electrical Engineer Civil Construction Engineer Instrumentation Engineer Project Process Engg / Production Engg Total no. of Positions

Electrical

Civil

Instrumentation

Chemical

40 70

For complete details w.r.t. eligibility criteria, selection process, etc. visit our website www.hindustanpetroleum.com

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HYDROCARBON PROCESSING MARCH 2012

I 25

HPIN CONSTRUCTION Toyo Engineering & Construction Sdn. Bhd. (Toyo-Malaysia), a Malaysian subsidiary of Toyo Engineering Corp., has a contract to build a 10,000-tpy bioethanol production facility for a joint-venture company between GlycosBio Asia Sdn. Bhd. and Malaysian Bio-Xchell Sdn. Bhd., located in Johor Bahru, Malaysia. GlycosBio Asia Sdn. Bhd. is a subsidiary of Glycos Biotechnologies Inc., an international biochemical company in the US, focused on commercializing renewable, high-value specialty chemicals. This plant is to produce industrial-grade bioethanol from crude glycerin. ToyoMalaysia will execute the project under a turnkey contract covering engineering, procurement, construction and commissioning (EPCC). The project is scheduled for completion in the second quarter of 2013. The plant will be developed in several phases, with a total production capacity rate up to 30,000 tpy being reached by 2014. A new process technology from UOP LLC, a Honeywell company, has been selected by National Refinery Ltd. (NRL) to maximize diesel and lubricant produc-

tion in Pakistan. UOP’s Uniflex processing technology was developed to help refiners processing the bottom of the barrel, the heaviest portions of a barrel of crude (also known as vacuum residue), into highervalue transportation fuels. NRL will use the Uniflex technology to upgrade its heavy residue into highvalue distillate products. Of particular value to NRL is the high yield of diesel from the Uniflex technology, which is said to be nearly double that of competing residue-conversion technologies. The technology will be integrated with UOP Unionfining hydroprocessing solutions to process distillates into high-quality diesel fuel and naphtha into high-quality feedstock used for gasoline production. Vacuum gasoil from the facility will also be converted to diesel and lube base oils using UOP’s Unicracking technology and fuels dewaxing technology provided by an alliance between Honeywell’s UOP and ExxonMobil Research & Engineering Co. (EMRE). The alliance, formed in 2011, brings together EMRE technology, for lube base oils production, with UOP hydroprocessing solutions, which pro-

Safety, Service, Quality

duce the high-quality feedstocks needed for lubricant production. The NRL facility, which is scheduled for startup in 2016, will produce 40,000 bpd of diesel fuel and 4,500 bpd of lube base oils. A Marubeni and Black & Veatch consortium has completed the Glow Phase 5 combined-cycle cogeneration plant in Rayong, Thailand, for the Glow Energy Public Co., Ltd. The low-emissions natural gas plant delivers a clean, reliable source of steam and electricity to Thailand’s growing industrial sectors, particularly the petrochemical industry. Black & Veatch served as the project’s technical manager. The company was responsible for engineering, balance-of-plant (all equipment except for the steam-turbine generator, combustion turbine and condenser) equipment procurement, construction management, startup and commissioning. Marubeni was the commercial manager, responsible for procurement of the power train and the construction subcontract. When operating in full condensing mode, the plant produces approximately

Imperial Crane Services, Inc. has been specializing in crane rental and sales for over 40 years. While growing and diversifying our ﬂeet we have focused on several areas of construction, including various work in the petrochemical industry.

Our recent new crane order of a multimillion dollar package of Sany crawler and rough terrain cranes, Tadano rough terrains and a 600 ton Liebherr was strategically purchased to handle our industrial heavy lift work throughout the country. The 600 ton LTM 1500-8.1 Liebherr makes us one of only a few companies with a crane of this capacity in North America.

Imperial has extensive experience in reﬁnery turnarounds and maintenance work. Our staff is skilled at managing multiple large scale projects with the ability to offer over 250 pieces of equipment, operator training, project management, cost estimation and lift coordination. • Hydraulic Truck Cranes 35 ton to 600 ton • Conventional Truck Cranes up to 300 ton • Crawler Cranes up to 352 ton • Rough Terrain Cranes 15 ton to 120 ton • Boom Trucks 10-50 tons with boom reach over 200’ • Industrial Elevators/Construction Hoists

2 Million Man Hours Without a Lost Time Accident (.61 EMR) Dealers in:

I MARCH 2012 HydrocarbonProcessing.com

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Bridgeview, IL: 708-598-2300 Grifﬁth, IN: 219-924-2900 St. Louis, MO: 1-888-HOIST IT Port Lavaca, TX: 1-888-HOIST IT www.imperialcrane.com

HPIN CONSTRUCTION 382 MWe of electricity. The plant can also produce 160 tph of steam in cogeneration mode, with an electricity output of 342 MW. Methanex Corp. plans to restart the second methanol plant at its Motunui site in New Zealand. The plant is expected to commence production in mid-2012 and will add up to 650,000 tpy of incremental capacity. Methanex has signed a 10-year agreement with New Zealand-based Todd Energy to supply natural gas for up to half of the 1.5 million-tpy capacity at the Motunui site. The plant has been idle since 2004 and the estimated future capital cost to restart is $60 million.

The facility will use several integrated UOP technologies to produce 1 million metric tpy of paraxylene and 500,000 metric tpy of benzene. UOP will provide engineering design, technology licensing, catalysts, adsorbents, equipment and technical service for the new units, which are expected to start up in 2014. The integrated aromatics complex will use the UOP Parex, Sulfolane, Tatoray, Isomar and CCR Platforming processes to produce high-purity paraxylene and benzene. The UOP Distillate Unionfining and Merox processes will be utilized to produce 2.6 million metric tpy of high-quality distillate products, including diesel and jet fuels.

UOP LLC, a Honeywell company, will provide technologies for producing key petrochemicals and high-quality diesel and jet fuel at Samsung Total Petrochemicals Co., Ltd.’s complex in Daesan, South Korea. The integrated aromatics complex will reportedly be the first to use UOP’s latest energy-efficient designs, allowing Samsung to reduce energy consumption by 33% and offering significant savings in greenhousegas emissions compared with earlier designs.

Toyo Engineering Corp. has a contract for a caprolactam plant from DSM Nanjing Chemical Co., Ltd. (DNCC), a joint-venture company between Royal DSM N.V. and SINOPEC Group. The project aims to install a 200,000-tpy caprolactam plant in Nanjing, China, in addition to the existing 200,000-tpy caprolactam plant. After the project completion, DNCC will reportedly become the largest caprolactam producer in the world with a 400,000-tpy capacity.

Toyo Engineering Corp. will engage in engineering, procurement services and construction management on a lump-sum basis. The client will invest approximately $300 million into the project, which is scheduled to be completed in the third quarter of 2013. Foster Wheeler AG’s Indian subsidiary, Foster Wheeler India Private Ltd., part of its Global Engineering and Construction Group, has been awarded a contract for a new LNG receiving terminal to be built in Ennore, Tamil Nadu, India. The contract was awarded by Indian Oil Corp., Ltd. (IOCL). Foster Wheeler’s scope of work includes basic design, front-end engineering design (FEED) and preparation of capital and operating cost estimates for the new LNG import, storage and regasification terminal. Foster Wheeler’s scope also includes marine studies to enable IOCL to finalize the jetty location, utilities and regasified LNG sendout facilities. The LNG terminal facilities, which will be designed to process 5 million tpy of LNG, are planned for completion by 2015–2016. HP

CONTRARY TO POPULAR BELIEF, NRI ENGINEERS DON’T READ MINDS. (But we’re amazingly good at letting clients pick our brains) Unlike other composite companies that just manufacture products, NRI has professional engineers who evaluate hundreds of elements to determine the precise composite solution to guarantee your pipeline’s integrity for years to come. We invite you to pick our brains and see why NRI’s innovative composite solutions and engineering know-how are recognized as the best in the industry.

561.683.6992 | neptuneresearch.com

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HYDROCARBON PROCESSING MARCH 2012

I 27

HPI CONSTRUCTION BOXSCORE UPDATE Company

City

Project

Ex Capacity Unit

Cost Status Yr Cmpl Licensor

KIMA Egyptian Refining Co

Aswan Cairo, Mostorod

Ammonia Hydrotreater, Naphtha

1200 tpd 22600 bpsd

Perdaman Inpex/Total E&P JV CSBP Ltd BP Zhuhai Chemical Co DSM Nanjing Chemical Co. BASF/SINOPEC/YPC PetroChina KPC/Sinopec/Total Essar Oil Ltd National Rfy Ltd Glow Energy Public Co Ltd

Collie Darwin Kwinana Guangdong Nanjing Nanjing Yunnan Zhanjiang, Donghai Island Vadinar Karachi Rayong

Urea LNG Ammonium Nitrate (3) PTA (3) Caprolactam (2) Ethylene Oxide Hydrotreater, Resid Ethylene Complex Amine Regeneration Unit Desulfurization, Diesel Cogeneration

6000 m-tpd 8 MMtpy None 1.25 m-tpy 200 tpy 330 Mtpy 75000 m-bpd 1 Mtpy 400 tpd 40 bpd 160 tpy

Suncor/Total E&P Canada Imperial Oil Ltd

Fort McMurray, Voyageur Upgrader (3) Kearl Oil Sands, Kearl Lake Bitumen

TO EX

200 bpd 110 bpd

ARMO Lukoil Neftochim Bourgas BASF KazMunaiGas Expl & Prod Rosneft CNPC Uz-Kor Gas Chemical LLC

Ballsh Burgas Ludwigshafen Atyrau Novokuibyshevsk Bagtyýarlyk Ustryat, Akchalak

Refinery Hydrocrack, Resid Toluene Diisocyanate (TDI) Extraction Lube Hydroprocessing Gas Processing (4) Propylene

None 2.2 MMtpy 300 m-tpy None 10 m-bpd 8 Bcmy 82 kty

Nasiriyah Saih Nihayda Sohar Doha Ras Laffan Ras Laffan Al Jubail Jazan Jubail Khursaniyah Yanbu Izmir Abu Dhabi

Hydrocracker Gas Compression Hydrocracker Acid Gas Removal Mono-Ethylene Glycol (MEG) Refinery EX Ethylene Vinyl Acetate Hydrocracker Calcium Chloride Gas Dehydration (2) Refinery Refinery TO Gas Treating

East Dubuque New Brunswick Sinclair

Ammonia Hydrocracker Hydrocracker (2)

Engineering

Constructor

KBR|Tecnimont GS E&C|Mitsui

Tecnimont GS E&C

AFRICA Egypt Egypt

U E

2014 2015

P P U E E C E U C U C

2014 2015 2014 2014 2013 2012 2014 2014 2012 2016 2012

U U

2016 2012

609

F E P P U U P

2012 2013 2014 2015 2014 2012 2015

m-bpd MW 266 bpd m-tpy m-tpy 6500 bpsd t/a 385 m-bpd None 300 MMscfd 501 Mbpd 1300 bpd 500 t/a 2200

E E E E S F E E E U E U E

2015 2013 2015 2013 2012 2013 2014 2014 2011 2014 2015 2013

U H C

2012

2044

KBR Technip|Axens KTI|ConocoPhillips

ASIA/PACIFIC Australia Australia Australia China China China China China India Pakistan Thailand

EX BY BY

2873 3400 20 300 1400 900 7.8

Chiyoda/JGC Corp/KBR Downer Clough Technip Toyo Engineering Corporation CLG UOP UOP

CLG Technip Black & Veatch

Essar Oil Ltd UOP Marubeni

CANADA Alberta Alberta

750

Jacobs CB&I

EUROPE Albania Bulgaria Germany Kazakhstan Russian Federation Turkmenistan Uzbekistan

1500 1300 1700

Axens

FW Technip

CLG

KBR

GS E&C

CLG

CLG GS E&C CB&I|CLG Technip|Chiyoda

Marubeni

MIDDLE EAST Iraq Oman Oman Qatar Qatar Qatar Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia Turkey UAE

SCOP PDOC Orpic Qatargas Shell Royal Dutch Qatargas GS E&C Aramco Services Co IDEA Pvt Ltd Saudi Aramco Saudi Aramco\Sinopec Socar\Turcas Enerji JV GASCO

50 15 96.8 10 1.5 292 200 106 120 400 214 35000

CLG

Shell|WorleyParsons

Technip GS E&C CLG Jacobs GS E&C Aramco Services Co|KBR Axens GS E&C

CLG CB&I

CLG KBR

GS E&C

GS E&C Axens GS E&C

UNITED STATES Illinois New Jersey Wyoming

Rentech, Inc. Irving Oil Ltd Sinclair Oil Corp

70 t/a 141 m-bpd 24 Mbpd

2013

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

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CORROSION CONTROL

SPECIALREPORT

Reduce CO2 in acid gas from amine-based TGTUs Improve furnace temperature and sulfur recovery B. SPOONER, Amine Experts, Kemah, Texas; and F. DERAKHSHAN, Sulphur Experts, Calgary, Alberta, Canada

Effects of CO2 on the sulfur plant. Slip is maximized

through the TGTU to prevent CO2 from re-entering the sulfur plant. There are several reasons why it is advantageous to minimize CO2 in the SRU feed, the most important involving the reaction furnace temperature, carbonyl sulfide (COS) and carbon disulfide (CS2 ) formation, and sulfur plant capacity, as outlined below. Reaction furnace temperature. CO2 does not burn in the reaction furnace and will, therefore, lower the furnace temperature. This will have negative effects on ammonia and aromatic destruction, which will, in turn, have consequences for the converter beds and the overall sulfur recovery. Maintaining reaction furnace temperature is critical in sustaining smooth and trouble-free operation of a sulfur plant. When processing sour water stripper gas, ammonia destruction is of primary concern, and reaction furnace temperatures of 2,250°F or higher are required to avoid ammonia salt formation and plugging of downstream process equipment. When processing acid gases containing aromatic hydrocarbons, reaction furnace temperatures of 1,920°F or higher are required to properly combust the benzene, toluene and xylene (BTX) components. Incomplete combustion of these compounds can lead to poisoning of the downstream catalyst bed. Finally, to maintain stable flame operation, a reaction furnace temperature of 1,700°F is required. Although some cracking can occur in the reaction furnace, CO2 is generally considered an inert compound, and it cools the

reaction furnace flame. As seen in Fig. 1, the dilution of an acid gas stream with increasing amounts of CO2 results in a rapid reduction of the reaction furnace temperature. Fig. 1 shows that, for a simple acid gas containing varying amounts of H2S and CO2 and 1% hydrocarbons, significant problems with maintaining the required reaction furnace temperature start to occur at very low concentrations of CO2 in the amine acid gas (AAG). With a CO2 content of less than 5% (based on this example), it becomes difficult to maintain the required reaction furnace temperature for good ammonia destruction. At a CO2 content of 30%, a breakthrough of benzene, toluene, ethylbenzene and xylenes (BTEX) could start to occur; and with a CO2 content of over 50%, problems with maintaining a reaction furnace flame become pronounced. Although not necessarily detrimental to Claus plant operation, higher CO2 content in the acid gas will require modifications to the operation of the reaction furnace in order to achieve and maintain the required operating temperature. Sometimes, these modifications come at a high capital investment cost or at the detriment of the overall efficiency. These items could include the addition of air and feed gas preheaters, oxygen enrichment, split-flow reaction furnace designs, installation of high intensity or other specialty burners, co-firing with fuel gas or acid gas enrichment. COS formation. Although the chemistry inside a reaction furnace is quite complex and chaotic, one certainty is that CO2 will 2,500 2,300

Thermodynamic temperature,°F Empirical temperature,°F

2,100 Temperature, °F

he purpose of an amine-based tail gas treating unit (TGTU) is to recycle any leftover sulfur components in the tail gas of a Claus sulfur recovery unit (SRU) to the front end of the plant, rather than incinerate them. The remaining sulfur components are converted to hydrogen sulfide (H2S) and removed from the gas using an amine solution. Carbon dioxide (CO2) is also present in tail gas streams and will be partially co-absorbed with H2S. This CO2 co-absorption should be minimized. Any CO2 removed in the TGTU will be recycled to the SRU front end along with the H2S, which has negative consequences for the reaction furnace and for the overall sulfur recovery. This article discusses the effect of CO2 on sulfur plants and how to minimize CO2 co-absorption or maximize CO2 slip through the TGTU’s amine absorber. A well-designed methyl diethanolamine (MDEA) TGTU should be able to achieve a minimum CO2 slip of 85%.

1,900 1,700 1,500 1,300 1,100 900 700 500 0

FIG. 1

30 40 50 60 70 80 CO2 in amine acid gas, mol% dry

100

Calculated adiabatic reaction furnace temperatures.

HYDROCARBON PROCESSING MARCH 2012

I 31

SPECIALREPORT

CORROSION CONTROL

partially “crack” in the furnace, resulting in the formation of CO, COS and (indirectly, through a drop in temperature) CS2. Fig. 2 shows the impact of CO2 on COS and CS2 formation rates, using the same simplified acid gas composition discussed previously. As shown in Fig. 2, the increasing CO2 concentration of the acid gas results in ever-increasing formation rates of CS2. CS2 formation has been shown to decrease with increasing temperatures, although it is not clear if less CS2 is formed at higher operating temperatures or if CS2 is formed and then quickly hydrolyzed. However, the addition of CO2 cools the reaction furnace temperatures, which results in increased CS2 formation. CS2 is especially unwanted since it binds up two sulfur molecules. COS and CS2 formation in the reaction furnace is important because, once formed, COS and CS2 do not participate in the modified Claus reaction. These compounds must, therefore, be converted or hydrolyzed back to H2S downstream of the furnace, either in the first Claus converter bed (operating at high temperatures, typically 600°F to 630°F) and/or utilizing special and often expensive catalyst; or they can be converted back to H2S in the TGTU hydrogenation reactor. Since amine-based TGTUs do not pick up COS and CS2 in the absorber, any unconverted COS and CS2 will result in increased sulfur emissions and reduced recovery efficiencies. Overall gas capacity. In a world of low-sulfur fuels, the demand for raising the existing processing capacity of refinery sulfur plants is becoming increasingly important. Since CO2 flows straight through the sulfur plant and does not participate in the Claus or modified Claus reactions, it takes up space and reduces the amount of sulfur-bearing gases that could otherwise be processed. A sulfur plant already operating at or slightly above its design capacity would quickly run into trouble with an increase in the CO2 content of the feed gas. Apart from the issues related to furnace temperature and increased COS and CS2 formation, operating a sulfur plant at higher-than-design throughput would create greater operating pressures that could result in: • Problems with air blowers not being able to deliver sufficient air at the higher operating pressures • Heat exchangers not being able to sufficiently cool process gases, resulting in additional sulfur vapor losses • Problems with mass velocities through condenser tubes, resulting in liquid sulfur carryover (both of which would require additional reducing gases in the TGTU hydrogenation bed). To address some of the reduced capacity issues, processing companies would have to consider technologies such as oxygen

Formation rate, % of inlet sulfur

35 COS formation rate CS2 formation rate

30 25 20 15 10 5 0 0

FIG. 2

20 30 40 50 60 70 80 Percent CO2 in amine acid gas, mol%, dry

100

COS and CS2 formation as a function of CO2 content in acid gas.

I MARCH 2012 HydrocarbonProcessing.com

enrichment or the construction of additional/larger sulfur plants, which would have a significant impact on the operating or capital budget of refineries. Effects of CO2 on TGTU quench water. High CO2 tail

gas streams serve to lower the pH of the quench water, which can decrease the strength of the protective iron sulfide (FeS) film. The water will, therefore, be darker in color and could cause fouling and plugging of the quench system. Ideally, the quench water has a pH of 7 to 8; however, in plants with high CO2 levels, the pH is typically between 6 and 7. Some plants try to correct this by regular caustic addition to the (partly) circulating quench water. Since a significant quantity of the quench water is not recycled, caustic injection is, at best, a temporary solution. Regular caustic injection can result in strong pH fluctuations that destabilize the protective FeS film. CO2 removal with MDEA. CO2 does not react directly with

the MDEA molecule; instead, it dissolves and reacts in the water portion of the solution: CO2 + H2O t H2CO3 (carbonic acid) H2CO3 t H+ + HCO3– (bicarbonate) H+ + R1R2R3N t R1R2R3NH+ CO2 + H2O + R1R2R3N t R1R2R3NH+ + HCO3– The reaction between CO2 and water (carbonic acid formation) is a “slow” step; it takes time to occur. Once the carbonic acid is formed, however, the MDEA reacts with it quite quickly, and the bond will not be broken again until the amine is regenerated. The removal of CO2 with a tertiary amine like MDEA is, therefore, kinetically limited by the reaction rate in the first step. H2S, on the other hand, reacts directly with the MDEA molecule: H2S + R1R2R3N t R1R2R3NH+ + HS– The reaction between H2S and the amine is a very fast or instantaneous reaction, which means that H2S removal is almost always equilibrium-limited. Each contact stage or tray in an absorber reaches the H2S equilibrium between gas and liquid. The difference in chemistry between H2S and CO2 removal is the key to understanding how H2S can be removed with a minimum amount of CO2. The best strategy to minimize CO2 removal with MDEA is to prevent the CO2/water reaction from occurring by: 1. Optimizing the amine temperature 2. Optimizing the amine strength 3. Optimizing the amine circulation rate 4. Optimizing the amine feed point of the absorber 5. Choosing a more selective solvent—e.g., formulated MDEA or sterically hindered amine. Amine temperature. As with most chemical reactions, the higher the temperature, the faster the CO2 reaction takes place. In a TGTU, lower lean amine temperature will minimize CO2 pickup because it reduces the reaction between CO2 and water. It should be noted that high amine temperatures (above 140°F) will also slip amounts of CO2. Unfortunately, at high temperatures, H2S will also be slipped due to equilibrium limitations. Therefore, operating the absorber at high temperatures cannot be used as an operating strategy. A low absorber temperature slows down the kinetics of the CO2 reaction. H2S removal occurs through a different mechanism

CORROSION CONTROL

H2S, ppm

400 360 320 280 240 200 160 120 80 40 0

H2S in vapor space CO2 in vapor space H2S higher ﬂow CO2 higher ﬂow

6.0 5.9 5.8 5.7 5.6 5.5 5.4 5.3 5.2 5.1 5.0

CO2, mol%

The contact time between the CO2 and amine depends on the internals of the absorber and the height at which the amine is injected. In a trayed absorber, contact time on each tray is determined by the weir height and the number of trays. If packing is used, contact time depends on the height and size of the packing. In either case, lowering the injection point of the lean amine will reduce the contact time and interfacial contact area between the gas and amine, thus increasing the amount of CO2 slip. Neither the weir height of trays nor the total height of packing is a parameter that can be adjusted during normal operation. Changes to either require shutting down the system and performing large-scale maintenance. This is why multiple feedpoints are built into the design of most TGTU absorbers. It should be noted that differences exist in the selectivity of packed and trayed towers. These differences are due to the hydraulics of the internals and the corresponding relationship with mass transfer. A trayed tower has the liquid phase, which is highly agitated; packed towers are opposite in that the liquid flows over the packing relatively smoothly. The gas flows are turbulent in both. Either type of internals can result in decent CO2 slip, but the choice between the two must be carefully considered and researched in the design stage. Since the net liquid holdup in a packed bed (1% to 6% of tower volume) is lower than in a trayed tower (8% to 12% of tower volume), the lower liquid hold-up can result in less interfacial contact area and, therefore, lower CO2 absorption. A TGTU absorber typically has three lean amine injection points: one at the top tray or top height of packing, and two more at successively lower intervals. Balance is achieved by injecting at a high enough point to remove all of the H2S necessary, but no higher. This procedure is best predicted first on a simulator,

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 Tray number

4,000 3,500 3,200 2,800 2,400 2,000 1,600 1,200 800 400 0

Effect of an increase in circulation rate on H2S and CO2 removal.

H2S in vapor space CO2 in vapor space

FIG. 4

1 2 3 4 5 6 7 8 9 10 11 12 13 14 15 16 17 18 Tray number

6.0 5.9 5.8 5.7 5.6 5.5 5.4 5.3 5.2 5.1 5.0

CO2, mol%

FIG. 3

H2S, ppm

that is much less affected by temperature. Therefore, the amine temperature should be kept as low as possible, although typically not lower than the inlet gas temperature. Since the gas is coming from the TGTU quench tower, it is saturated with water. An amine temperature lower than the gas temperature would condense water and dilute the amine solvent. It should be noted that the normal temperature guideline for amine systems—maintain 5°F to 10°F temperature difference between gas and amine—is not applicable for amine-based TGTUs. Normal operation is to maintain the lean amine temperature at or within 1°F to 2°F of the inlet gas. An ideal temperature range for both the amine and the gas is 90°F to 100°F. Amine strength. Since water is one of the reactants, the more water there is in the absorber, the more CO2 will be absorbed into the amine solution. MDEA can be successfully operated at up to 50% strength, with the other 50% consisting of water. Above this strength, the viscosity of the solution becomes too high and will negatively affect the mass transfer of H2S into the amine. The CO2 hydrolysis should be minimized by limiting the water content of the MDEA solution. Solvent strength should be maintained between 45% and 50% while utilizing a proper filtration program to ensure that the viscosity of the amine does not rise too high. Minimizing SO2 breakthrough from the Shell Claus Offgas Treatment (SCOT) or SCOT-type cobalt-molybdenum (CoMo) reactor will reduce the heat-stable amine salt buildup in the system and minimize the chance of a viscosity increase due to excess heat-stable salt formation. Amine circulation rate. The CO2 reaction with water (to form bicarbonate) is a kinetically limited reaction. This means that CO2 builds up at the gas/liquid interface and only reacts with the water and amine as it slowly diffuses to the bulk of the amine solution. In a trayed tower, higher circulation rates increase the height of the liquid on each tray, as the weir creates a flow obstruction. The amine “stacks up” against it and provides a larger surface area for CO2 absorption. In a packed tower, greater amine flows create higher holdup of amine into each section of packing. The more liquid in the tower, the higher the CO2 removal will be, as there is more gas/liquid interface for the CO2 to be absorbed. Overcirculation of the amine is the single largest contributor to poor CO2 slip in TGTUs. Furthermore, CO2 increases the loading of the amine, which takes up valuable acid gas holding capacity, especially in lowpressure TGTU applications. Under certain conditions, it is possible that CO2 will force the amine to release previously absorbed H2S. Optimizing (i.e., reducing) the amine circulation rate will always result in a decrease in H2S and an increase in CO2 in the treated gas, which is the goal. In Fig. 3, the effects of circulation rate are shown. As the amine rate is decreased from 200 gallons per minute (gpm) to 150 gpm, the CO2 in the treated gas increases by 1,000 ppm, whereas the H2S decreases from 45 ppm to 40 ppm. The circulation rate should be initially targeted for a rich H2S loading of 0.05 mol/mol, and slowly decreased with a final target loading of 0.1 mol/mol. This should be reinforced by the use of a reliable amine plant simulator. Amine feedpoint into absorber. TGTU absorbers often have multiple inlet points for the lean amine solution. Several inlet points are normal and can be used to increase or decrease the interfacial contact area between the gas and the amine solution. A shorter contact time (less interfacial contact area) will result in less CO2 absorption.

SPECIALREPORT

Simulated acid gas removal tray-by-tray in a TGTU absorber. HYDROCARBON PROCESSING MARCH 2012

I 33

SPECIALREPORT

CORROSION CONTROL

as shown in Fig. 4. Using this type of chart, it is apparent where appropriate injection points into the absorber exist, depending on the required H2S specification. If possible, the amine injection should be lowered by one feedpoint, and the treated gas H2S content should be measured. If it is acceptable and simulations agree, the amine injection can be lowered to another feedpoint. The main concern here is the H2S content of the treated gas.

By lowering the H2S lean loading, the H2S in the treated gas will also drop. This could enable the lean amine to be introduced into a lower feedpoint in the absorber, thus increasing the CO2 slip, as described above. Normally, however, the aim of the improved regeneration is to obtain a lower H2S level in the treated gas. TGTUs are, therefore, often designed with more trays in the regenerator than what is normally seen in amine units; this allows for deep H2S stripping. HP

Choosing a more selective solvent. The use of an amine

with high selectivity can engender a number of benefits for the TGTU, as discussed in the following two options. Hindered amine. The most selective solvent for TGTU applications is a hindered amine, which is normally a secondary amine with a bulky group that hinders the direct reaction with CO2. These molecules combine the low CO2 reaction rate of a tertiary amine with the base strength of a secondary amine. The high base strength is particularly useful at the low pressure of a TGTU absorber because it makes it possible to have a significantly higher rich amine loading than with MDEA. The higher loading allows for a reduction in amine circulation rate, which further improves the selectivity. Formulated MDEA. Enhanced MDEA formulations for TGTUs have been available for 25 years. These are normally the solvents that contain “pH suppressants” such as phosphoric acid. The reduction in pH allows for easier and deeper regeneration of the amine, especially for H2S. The same effect of improved regeneration is achieved in MDEA solutions containing between 0.5 wt% and 1.0 wt% heat-stable salts.

I MARCH 2012 HydrocarbonProcessing.com

Ben Spooner, a senior process engineer, has been working in the amine industry as an operator and engineer since 1998. He joined Amine Experts in 2003 and has worked in over 25 countries and on hundreds of amine systems, providing expert assistance and advice regarding plant operations, troubleshooting, optimization and operator training. Mr. Spooner is one of the primary speakers at Amine Experts’ world-recognized Amine Treating Seminar, which has been presented in dozens of locations around the globe. He holds a BSc degree in petroleum engineering from the University of Alberta.

Farsin Derakhshan is a professional engineer with over 16 years of experience. He joined Sulphur Experts in 1996, initially working out of the German office, and since then he has been directly involved in all aspects of Sulphur Experts’ process engineering consulting work. Now residing in Canada, Mr. Derakhshan is an experienced and well-traveled sulfur plant engineer, providing expert advice and consulting services to clients around the world. His specialty areas include sulfur plant troubleshooting and process optimization. Mr. Derakhshan is also Sulphur Experts’ regional engineer for Europe and the Middle East. He is technically responsible for all projects in those regions, and he is also a guest speaker at the internationally recognized Sulphur Recovery seminars. He holds a BSc degree in mechanical engineering from the University of Calgary.

Select 159 at www.HydrocarbonProcessing.com/RS

CORROSION CONTROL

SPECIALREPORT

Consider a new monitoring system to prevent corrosion Innovative, continuous supervising method collects real-time data on key asset health P. COLLINS, CEO of Permasense Ltd., UK

combination of aging plants, greater fluid corrosiveness and tightening of health, safety, security and environment (HSSE) requirements has made corrosion management a key consideration for refinery operators. The prevention of corrosion/erosion through live monitoring provides asset and integrity managers with a real-time picture of how their facility is coping with the high demands placed upon it by corrosive fluids. This information can assist in risk management and auditing. Continuous measurement presents a step change in the level of corrosion rates that can be determined and the accuracy of that determination. Plant integrity. Steel pipework and

vessels are always at risk of corrosion or erosion. Unless monitored, there is a risk of failure, which may impact the safety of workers and the environment. The financial costs of operational interruption, repairs and reputational damage must also be considered. As oil and gas operators produce and process ever more corrosive or erosive hydrocarbon streams, the demands on plant metallurgy steadily increase. Permanently installed sensor systems can deliver a continuous picture of asset condition over time, at a comparable cost to that of a single manual inspection. This picture can be correlated with process conditions that may be causing corrosion or erosion, and strategies to minimize corrosion, such as inhibitor use. With such knowledge, the asset manager can move beyond merely knowing whether corrosion or erosion is occurring, to understanding why and at what rate. This understanding enables operators to make betterinformed decisions.

Need for continuous monitoring.

There are various established techniques for the periodic assessment of pipe and vessel integrity. The drivers of corrosion and erosion—process conditions, crude constituents and abrasive solids—and the inhibitors to hold corrosion rates in check are familiar. Periodic inspections do not, however, deliver continuous pipework condition data that can be correlated with either corrosion drivers or inhibitor use to understand the impact of process decisions and the inhibitor usage on plant integrity. Manual acquisition of ultrasonic wall thickness data is also frequently associated with repeatability limitations and data-logging errors. Permanently installed sensor systems, on the other hand, deliver continuous highquality data. The ultrasonic sensors can be installed on pipes and vessels operating at up to 600°C (1,100°F). These sensors have also been certified as intrinsically safe for use in most hazardous environments. The

FIG. 1

system has been proven in operation over a number of years in refinery environments, and more recently in upstream facilities. Continuous monitoring installation data can validate that, when corrosion is occurring, it is often an intermittent process rather than a continuous event. It is in such cases that it is particularly valuable to be able to correlate thickness data over time with process and/or inhibitor parameters. Moreover, the data highlights which prevention or mitigation strategies are most effective. System design. At the core of the continuous monitoring system is an ultrasonic sensor mounted on stainless steel (SS) waveguides. The waveguides isolate the sensor electronics from extreme temperatures and guide the ultrasonic signals to the pipe wall and back without excessive signal degradation or distortion. The system can monitor pipe wall thicknesses in the range

Wireless communication of the continuous corrosion monitoring system.

HYDROCARBON PROCESSING MARCH 2012

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SPECIALREPORT

CORROSION CONTROL

of 3 mm to 40 mm (1⁄8 in. to 1½ in.) and can be applied on a wide range of steels and other alloys. Frequent measurement of wall thickness allows for metal loss detection at the level of 10s of microns. Each sensor is equipped with a radio, and communicates with other sensors and a gateway (base station) within a 50-m (55-yd) range. The sensors form a mesh or wireless network that does not require previous installation of wireless network infrastructure (Fig. 1). Each sensor radio can also act as a relay, or repeater, enabling the network to span hundreds of meters from the gateway. The data is channeled via the gateway to a database on a connected computer. If, as is the typical case, this computer is networked, then browser-based visualization software enables the corrosion/inspection engineers to view the data at their desks.

The data can also be exported in any of the file formats required by the various process monitoring applications, enabling seamless transfer and read-in to those packages and, thus, correlation with the process data at the sensor location. The principles of the system were developed by the world-leading nondestructive testing research group at the Imperial College London, led by Professor Peter Cawley. It was refined and proven over several years of collaboration with BP refineries. The experience gained in this collaboration helped produce the robustness required for harsh refining environments. The system was conceived from the outset to be costeffective for large-scale deployment. Cost-effective for large-scale deployment. The sensors are battery

which minimizes the installation costs and imposes fewer restrictions for remote areas and for large-scale deployments. The sensor is secured on the pipe/vessel by means of two studs that are welded onto the pipe. For pipe-wall temperatures below 100°C, the studs can also be welded onto girth clamps, which are themselves mounted on the pipe. Stud mounting allows for dry coupling; no couplant is required between the waveguide tip and the pipe wall. This, together with multiyear battery life, eliminates the need for expensive maintenance access between turnarounds. Stud-based mounting also enables geometric flexibility and reduces installation time to just minutes. A two-person installation team can typically install 50 sensors per day.

powered. Thus, no cabling is required, Robust wireless communication.

FIG. 2

FIG. 3

Measured wall thickness for sensors installed on one carbon-cast steel U-bend in the field.

Monitoring data of the carbon steel U-bend at BP’s Gelsenkirchen refinery.

I MARCH 2012 HydrocarbonProcessing.com

The sensor has been designed using highgrade materials to allow for many years of continuous operation. A number of systems have been in uninterrupted operation for three years. To ensure that the system performs in the event of a blockage of an individual pathway or the loss of a sensor, there are multiple pathways for data transmission through the mesh back to the gateway (Fig. 1), which guarantees data retrieval. The gateway channels data transmitted from all the sensors located in the network. Typically, wall thickness measurements are sent every 12 hours. This interval can be changed at any time for any sensor, to as little as a few minutes if necessary, depending on the monitoring or metal loss determination requirement at that location. Data is stored in the computer database to guarantee security. This also allows the user to view a full history of data readings, and build a clearer picture of corrosion and erosion rates. Applications. The system has a wide

range of applications in the hydrocarbon processing industry. At present, nearly 20 refineries now use this corrosion monitoring system and it is in use on virtually all crude unit lines, air coolers, furnaces, heat exchangers, pumps, amine units, cokers and cracking units. Pipe materials include carbon, chrome and stainless steel. Typical locations for sensor installation are on elbows, which are known as thin spots, and areas of particular turbulence. Older units, particularly those operating outside of design specifications, are worthy of attention.

CORROSION CONTROL The system allows facility operators to monitor locations continuously without the repeated cost of access. By correlating metal loss data with process data (composition, hold-up, temperature), a true understanding can be gained of what changes in parameters are driving corrosion and erosion processes. This understanding is enabling operators to make better-informed decisions about changes according to these parameters to minimize the impact of corrosion on their plant. Furthermore, users are now optimizing their inhibitor and biocide use, by level and location, based on insights gained from the data. Continuous monitoring on near-endof-life lines enables turnarounds to be scheduled with much greater confidence. In a recent example, a system installed on a line with an expected remaining life of 12 months enabled line replacement to be postponed by a very valuable six months. Plus, the used sensors were recovered for re-installation elsewhere. Inspector safety. In plants with

aggressive rates of corrosion, particularly where corrosion is intermittent and the remaining life is uncertain, frequent manual inspection is common. Where operating temperatures are sufficiently high and a shutdown is necessary for safety reasons to enable manual inspection, the lost production can come at a high cost. Some locations in a facility can be hard to reach; thus, technicians incur safety risks in gaining access. Where high pipework and vessel temperatures are involved, ensuring technician safety during manual inspection becomes even more challenging. Permanently installed systems reduce the safety risks associated with collecting plant condition data. In several chemical production facilities, corrosion monitoring system users have also been able to eliminate the periodic shutdowns that they had previously required to enable operator access. The installed systems are also now delivering data where inspector availability is limited or where access is difficult for environmental reasons, such as in Arctic locations.

ued safe operation (Fig. 2). Since the high temperature prevented accurate manual ultrasonic wall-thickness measurement, and would have exposed inspectors to significant hazard, the continuous monitoring system was installed, as shown in Fig. 3. This secured operation with confidence until a turnaround. The system has been delivering reliable measurement data for three years. New monitoring method. Operating companies using the continuous corrosion monitoring solution have a more accurate and timely understanding of the corrosion and erosion rates occurring in their facilities. Where inhibitors are in use, the system is giving a greater understanding of their effectiveness. The real-time data allows potential corrosion hotspots to be remotely monitored, at time intervals of the operator’s choosing. This insight allows asset managers to make more informed decisions, to the benefit of plant integrity, safety and operating costs. The system has been tried and tested in some of the most inhospitable environments, and it operates at pipework tem-

SPECIALREPORT

peratures from –30°C to 600°C (–20°F to 1,100°F). It allows operators the freedom to choose monitoring locations irrespective of how inaccessible they are, thanks to the use of ultrasonic sensors and wireless networks for data retrieval. Having already been installed for a number of years in BP refineries across the world, the system has now been adopted by other super-major and privately held refinery operators in the US, Germany and Canada. This technology is making a real difference in an industry facing new challenges every day. HP Peter Collins is the CEO of Permasense. Dr. Collins joined the Permasense board in 2010, and he is responsible for the overall development of the company. Already an experienced entrepreneur, he has held board-level technical and operational roles in public and private companies with $100 million+ revenues. His previous roles include operations director at Sondex plc, which specializes in the engineering and manufacture of directional drilling and formation evaluation systems and wireline tools for the production of oil and gas. He was earlier technical director at PII Ltd. and a manager at management consultancy Arthur D. Little Ltd. Dr. Collins holds a PhD in computational fluid dynamics from Imperial College London, a BE degree in mechanical engineering from the University College Dublin, and an MBA from INSEAD.

Gelsenkirchen experience. Corro-

sion monitoring was conducted on cast carbon steel U-bends with a wall thickness of approximately 25 mm (1 in.), operating at 380°C (720°F) in the Gelsenkirchen refinery operated by BP to ensure continSelect 160 at www.HydrocarbonProcessing.com/RS

Turnaround Welding Services

“These guys are doggone good!” I’m proud to be the mascot for the workers at Turnaround Welding Services, because they are the best in the business. Why? It’s because they tackle a planned outing just like they tackle an emergency. They put their hearts, their souls, their backs and their expertise to work and their reputations on the line…every day. They don’t just do the big jobs, either. They handle jobs that may only take ten man-hours as well as the ones that require over 250,000 man-hours. Another reason I’m proud to be with Turnaround Welding is that our guys aren’t spoiled brats. They are trained to perform multiple tasks. They might weld, ﬁt and rig the same piece of equipment and not even ask for a helper to carry their leads. They’re so good that they don’t have to brag––even though their weld reduction rate is low and their productivity is high. And their safety record, well, it’s about as good as it gets. Yeah, these guys are good, all right! They deliver their services with the tenacity of a bulldog––and I know all about that.

1.225.686.7101 or visit www.turnaroundweldingservices.com Select 78 at www.HydrocarbonProcessing.com/RS

CORROSION CONTROL

SPECIALREPORT

Avoid stress corrosion cracking of stainless steel This case history investigates equipment failure in a glycol unit A. D. JAIN, Reliance Ports and Terminals Ltd., Navi Mumbai, India

n hydrocarbon processing industry (HPI) facilities, austenitic stainless steel (SS) is used for several reasons, including: • Avoiding iron pickup and maintaining product purity • Replacing carbon steel in operating services in which carbon steel is unsuitable, such as high temperature, corrosive fluids, highly toxic materials (HTM), etc. However, austenitic SS is susceptible to stress-corrosion cracking (SCC), mainly through trans-granular stress-corrosion cracking (TGSCC), such as chloride-containing aqueous media. The presented case history investigates the SCC failure of SS in a petrochemical unit manufacturing high-purity monoethylene glycol (MEG). The MEG refining sections are composed of insulated columns with stiffening rings to provide resistance against buckling under vacuum conditions. Process equipment is exposed to appreciable levels of chlorides due to: a) Rainwater, in cases of atmosphere where the salinity is greater than normal b) High-temperature insulation materials that carry leachable chlorides (more than 50 ppm). Also, the heat exchanger connected to the column showed cracks on the dome. In such cases, equipment and piping made of austenitic SS is more prone to SCC. These types of cracks initiate at the heat-affected zone (HAZ) of welds between the stiffening rings and columns. This article details the failures, along with the related inspection and testing methods. A good approach to delay, and, if possible, avoid SCC is to mitigate its initiation. Several prevention methods are discussed, along with a comparison review between different SSs.

Corrosion never sleeps. SCC occurs due to a process involving the conjoint corrosion and straining of a metal due to residual or applied stresses.1 SCC is an insidious form of corrosion; it produces a marked loss of mechanical strength with little metal loss. The damage is not always obvious from casual inspections. Stress-corrosion cracks can trigger mechanically fast fractures and catastrophic failure of components and structures. The occurrence of SCC depends on the simultaneous achievement of three requirements, as shown in Fig. 1: • A susceptible material • An environment that fosters SCC of that material • Sufficient tensile stress to induce SCC. TGSC cracking. This process involves an active path dissolu-

tion mechanism in which the accelerated corrosion occurs along a path of higher-than-normal corrosion susceptibility, with the bulk of the material being passive.1 Fig. 2 shows the schematic plan for TGSCC mechanism. Initially, pitting corrosion occurs in the presence of a chloride environment. This process can occur in the absence of stress, giving rise to small pits that form a homogeneous brownish layer. The effect of the applied stress is mainly due to cracks that form from the pits over a long period and are accompanied by process temperature upsets, chloride attack, etc. These conditions foster easier diffusion of corrosion products away from the crack tip and allow the crack tip to corrode faster.1

Corrosion: Chloride attack followed by small, localized pit formation

Speciﬁc corrodent

Stress corrosion cracks FIG. 1

Mechanical rupture: Under action of residual tensile stress

Stress

System diagram illustrating the minimum requirements for SCC.2

Crack growth: Corrosion + rupture FIG. 2

Result: SCC

Progression and stages of SCC.3

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SPECIALREPORT

CORROSION CONTROL

Concentrator section SCC

11% glycol and rest water

68% glycol and 32% water

E26

Evaporation section

C12

C13

98% glycol and water vapor To reﬁning section

FIG. 3

Flow diagram of the evaporation and concentration section of the MEG unit.

FIG. 4

a) Visual of the MEG column shows cracks at HAZ of weld joint between stiffening ring and column; b) nondestructive testing with soap bubbles identifies leak locations.

FIG. 5

DP test on Van Stone flange of the heat exchanger reveals stress-corrosion cracks.

FIG. 6

DP test on the heat exchanger dome reveals stresscorrosion cracks.

TABLE 1. Operating condition of final concentrator column C13

TABLE 2. Operating condition of exchanger E26

Temperature: Top: 77°C Bottom: 178°C

Pressure: 421 mbar tube side

Pressure: Top: 421 mbar, Bottom: 527 mbar (vacuum condition)

Service fluid: MEG (0.99%) + water vapor (99%) tube side and cooling water shell side

Service fluid: MEG (0.99%) + water vapor (99%)

Equipment failures in MEG section. Our focus will

center on two pieces of equipment that succumbed to SCC: 1) the final concentrator column C13, and 2) the final concentrator condenser E26 and connected piping. Overview of MEG section. The ethylene glycol (EG) concentrator section receives the concentrated glycol stream from the glycol evaporation section, which contains glycol and water in a 68:32 ratio. The glycol concentrator section consists of two concentrator columns. The pre-concentrator and final concentrator are the trayed columns that remove the remaining water from the glycol. The refining section receives 98 wt% glycol from the concentrator section and produces a high-purity polyester-grade MEG product, as shown in Fig. 3. The entire unit was commissioned in 1989. After 20 years of safe operation, a sudden equipment failure was noticed in this section. Final concentrator column C13 failure. The unit failure came in the form of major cracks at the HAZ of the weld joint between the stiffening ring and the shell of the column. During normal running conditions, no abnormalities were noticed. However, when the plant was started up after a maintenance shutdown, frequent vacuum breaks were noticed in column C13. The plant was again shut down for inspection to find the root cause of the abnormality. The insulation of the entire column was removed; at the stiffening ring location, significant salt deposition was detected. The protective SS foil wrapping was also not present 40

I MARCH 2012 HydrocarbonProcessing.com

Temperature: Inlet: 77°C, outlet: 55°C tube side

inside the insulation at the stiffening ring location. The column was pneumatically tested via a soap solution, and through-andthrough leakage was observed, as shown in Fig. 4. It was concluded that these cracks were already present before shutdown but opened due to the temperature fluctuations during shutdown/startup cycles. The cracks would have opened further if the shutdown had not been taken. Final concentrator condenser failure at Van Stone flanges. During the shutdown, the heat exchanger was hydrotested. Pressure drop accompanied by water leakage was noticed at the exchanger dome. After removing the insulation, dye penetrate (DP) testing showed major SCC cracks of the dome at the stubend location. The connected pipeline loop having a Van Stone flange was also checked by DP testing. Figs. 5 and 6 show DP testing on the Van Stone flanges of the piping section and the exchanger, respectively. Inspection and testing. After failure, detailed inspection and testing were completed for the column at the stiffening location and the exchanger dome at the Van Stone flange. Visual inspection. The surface of the final concentrator column can be divided into several sections, as shown in Figs. 7–9. • In Fig. 7, the surface of the column was not affected by SCC away from the HAZ region of the weld joint between the stiffening ring and column shell.

CORROSION CONTROL

SPECIALREPORT

FIG. 9 FIG. 7

FIG. 10

Silver surface indicates no SCC present.

FIG. 8

Brownish spots shows early signs of surface pitting.

Radiograph taken at HAZ location.

â&#x20AC;˘ In Fig. 8, the surface is closer to the HAZ region of the weld joint, but away from the region where water logging and salt concentration occurred. The surface has minor pits, which could be the initial stage of an active path dissolution mechanism. â&#x20AC;˘ In Fig. 9, in the HAZ region, continuous water logging occurred due to water ingress through the insulation cladding during the monsoon seasons. Radiographic examination. Radiographic examination via a double-wall single image technique was done at the HAZ region of the weld joint. Fig. 10 shows the radiographic image. A continuous chain of branched cracks can be seen. DP examination. A DP examination of the heat exchanger and connected piping section with the Van Stone flanges from inside and outside was done to detect the origin of the cracks, as shown in Fig. 5. The DP test of the heat exchanger, as shown in Fig. 6, identified major indications at the outside surface and only a few indications at the inside surface. Indications from the inside are only attributed to through-and-through cracks that propagated from cracks initiating from the outside surface. This confirms that the cracks had an external origin. Metallographic examination. Metallographic examination was done as per ASTM A262 Practice A, where cracks were noticed in the visual examination.5 As can be seen in the micrographs from Fig. 11, TGSC cracks were found. Chemical analysis of the insulation. Chemical analysis of the used insulation material from the MEG column was carried out as per ASTM C871.6 For comparison purposes, new (unused) refined mineral wool insulation and light resin bonded mineral wool insulation were also chemically tested as per the same standard. The laboratory test report is given in Table 3. The used insulation of column C-13 showed a 78-ppm level of chloride and approximately 400-ppm levels of sodium and silicates (Na + SiO3). As as per Fig. 12, it should have about a 600ppm level of Na + SiO3 to act as an inhibitor and to counter the presence of chlorides. The new refined mineral wool insulation

FIG. 11

Blackish spots indicate water logging and water ingress on the column surface.

The photos show TGSCC in highly sensitized austenitic grains at the HAZ region near the weld joint of the stiffening ring with the shell of the column; taken at 200X magnification.

is an inhibiting insulation as per ASTM C795.7 The chloride-tosodium and silicates levels are acceptable as per the requirement given in Fig. 12. This insulation type is expected to prevent SCC. Probable reasons for failure. The service fluid (MEG +

boiler feedwater) does not have any detrimental effects on SS 304 type material. For years, the austenitic SS 304 grade material was used in MEG. No adverse effects by MEG on SS 304 type material were reported. During an internal inspection of the column, no abnormalities were noticed. Accordingly, the probable cause for the internal degradation can be ruled out. Inspection and test results also indicate that the cracks were from an external origin due to a chloride attack on the susceptible SS 304L materials. Failure of final concentrator column. In vacuum columns, the stiffening ring is used to avoid buckling from internal and external pressure differences, as shown in Fig. 12. The stiffening ring is welded circumferentially to the column shell. The HAZ region of the welded stiffening ring, if not stress relieved by solution annealing, can remain sensitized. In saline atmospheres, rainwater contains chlorides. If rainwater permeates through the insulation, it can accumulate on the stiffening ring. Over time, the water can concentrate and form chloride salts, which can further concentrate and form solid chloride salts. Physical salt deposits were noticed after removing the insulation. The stiffening ring provides ample space for water logging, as shown in Fig. 13. The mineral wool insulation material carries leachable chlorides. The presence of chlorides at sensitized regions can form small pits, and the induced tensile stresses can trigger cracks from such small surface openings. The conditions finally result in major SCC at the stiffening ring locations. Minor surface cracks were also observed at the HAZ location of the RF Pad. HYDROCARBON PROCESSING MARCH 2012

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CORROSION CONTROL

Stiffening ring design considerations. Fig. 14 shows several acceptable options for stiffening ring attachments to the column shell as per UG-30 of ASME Sec. VIII Div. 1.8 The attachment (Fig. 14 b) was used for column C13. Instead of (b), using attachment (d) or (e) designs could avoid water logging. Acceptable stiffening ring welded attachments are shown in Fig.

Susceptible locations for SCC in piping. Insulated pip-

Y axis Used C-13 insulation X axis: 600 ppm Y axis: 78 ppm LRB X axis: 200 ppm Y axis: 60 ppm Reﬁned insulation X axis: 70 ppm Y axis: 26 ppm

10,000

Chloride, ppm

15 as per clause UG-30 of ASME Sec. VIII Div. 1. The attachment (c) design of Fig. 15 was used for column C13. Inline intermittent (a) or staggered intermittent (b) welding of the stiffening ring with the column shell can also be used as per the design code, which will not only prevent water logging but also decrease the heat input requirement during the welding process.

Unacceptable

1,000

Acceptable

100 78 60 26 10 10

FIG. 12

10,000 10 100 200 600 1,000 Na + SiO3, ppm

100,000 X axis

Lab results of used insulation for chloride, salts and silicates concentrations; acceptability criterion is per ASTM C 795.

ing and equipment are more prone to SCC, as the insulation carries leachable chlorides. Fig. 16 shows the typical piping locations where water ingress and logging are a common phenomenon. Over time, the sandwiched water between the insulation and metal surface evaporates and concentrates the chloride solution. SCC is prone at the locations where saltwater accumulation and induced stresses are present. The origin of induced stresses could be due to fabrication and service factors. For example, weld joints can induce stresses at HAZ locations, which are also sensitized and where the possibility of developing initial pits is greater. Similarly, the Van Stone flanges or lap-joint flanges have induced stresses at the stub-end location, as shown in Fig. 17. Failure in exchanger flange. At location “A,” as shown in Fig. 17, there are three requirements for SCC that must be simultaneously present, and these include:

TABLE 3. Laboratory testing of column insulation materials per ASTM C871 Water logging and accumulation of chlorides

Column shell ASTM A240 TP 304

Stiffening ring IS 226/IS 2,062 Major cracks at HAZ of weld joint between stiffening ring and shell

Minor surface cracks at HAZ of RF pad

FIG. 13

Section of MEG column water logging location at stiffening ring.

Used column C13 insulation

New light resin bonded insulation

Refined mineral wool, Spec. ASTM C705

Moisture @ 110°C, %

0.63

0.67

0.84

Al, ppm

13.9

24.7

Ca, ppm

237

294

408

K, ppm

16.9

Fe, ppm

5.3

NIL

Mg, ppm

Na, ppm

255

250

Si, ppm

51.9

Parameters

Chlorides, ppm Silicates as SiSO3, ppm

165

111

141

Remarks: Chlorides are higher in the used and new light resin bonded insulation samples.

Stiffener tw

Shell

FIG. 14

2-in. (50-mm) min.

tw t

(e)

Acceptable stiffening ring attachments with the column shell as per clause UG-30 of ASME Sec. VIII Div. 1.

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Continuous full penetration weld

w (d)

2-in. (50-mm) min.

24t max.

w (b)

(a)

2-in. (50-mm) min.

tw w

In-line intermittent weld (a) FIG. 15

Staggered intermittent weld (b)

Continuous ﬁllet weld on one side, intermittent other side (c)

Acceptable stiffening ring attachment welding details as per clause UG-30 of ASME Sec. VIII Div. 1.

CORROSION CONTROL 1. Induced tensile stresses during the forming process are the maximum, as the radius of curvature is the minimum at this point 2. Chloride, in the form of salts, is concentrated in the space between the flange and the stub end, and the source is rainwater followed by evaporation and/or leachable chlorides from insulation material 3. Susceptible material is SS 304. Case study 1. This case study investigates the SCC behavior of SS 304L and SS 304LN in 5% HCl solution.9 The experiment was conducted as per ASTM G36. In the study findings, the U-bend samples of SS 304LN were observed to fail through the ductile mode, and they required a high load to fracture, as shown in Table 4. Conversely, brittle failure occurred in SS 304L samples at a lower load. Result: SS 304L is more susceptible to SCC than SS 304LN. Based on the inspection, testing and analysis carried out, the equipment failure on the MEG unit resulted from these conditions and processes: 1. Water ingressed to the SS surface through insulation stemming from gaps in the cladding joints. 2. SCC occured due to chloride attack at the sensitized HAZ location of the weld joint between the stiffening ring and the column shell. 3. Severe SCC at the stiffening ring, as compared to minor cracks at the RF pad, indicated that the equipment locations were externally water logged and more prone to SCC. 4. SCC also occurred at the stub-end location of Van Stone flanges due to the simultaneous action of chlorides and a high level of induced stresses. 5. SCC was of an external origin, as confirmed by inspection and testing. 6. Metallographic examination revealed that the cracks were of a branched TGSCC type. 7. Chemical analysis showed that the leachable chlorides from column C-13 “used insulation” could have led to SCC in SS materials and this was also not acceptable as per ASTM C795. However, the new refined mineral wool was acceptable.

SPECIALREPORT

Mitigation and remedial measures. Several recom-

mended actions were taken to mitigate cracking of the stiffening ring for the MEG column: 1. Cracks that are branched cannot be repaired by grinding and welding. These cracks can grow further during the welding process. A new column could not be installed at the time of the present shutdown. The alternative remedial action was a coldcompound repair. The cracked portion around the stiffening ring was strengthened by applying a cold compound followed by fiberglass cloth impregnation with a water-activated resin. Strengthening was done on both the stiffening ring and the RF pad sections of the MEG column. 2. Holes were drilled in the stiffening ring to avoid further water-logging. 3. New refined-mineral wool insulation, acceptable as per ASTM C 795, was selected for re-insulation purposes. 4. Cladding joints were sealed using putty. 5. In designing the new column, instead of a Van Stone flange, weld neck or slip on the flange was selected for the nozzles and connected pipelines. 6. A change in alloy composition from SS 304L to SS 304LN was recommended for the new MEG column design. Cracks at Van Stone flanges of heat exchanger and connecting piping. Several design and remedial actions were recommended for the heat exchanger and connecting piping: 1. In the case of the heat exchanger, the cracked portion was removed by cutting. A new lap-joint flange was welded. 2. The lap-joint location was sealed by a cold compound to avoid ingress of rainwater. 3. New refined-mineral wool insulation was used. Lessons learned. It is important to avoid water ingress into the insulation by sealing the cladding joints. Van Stone or lapjoint flanges are not suitable for SS piping systems installed on critical services. When rainwater ingress is an issue, the stiffening ring design should include holes or intermittent welding with the column shell to drain out the collected water.

TABLE 4. SCC behavior for SS 304L and SS 304 LN in 5% HCl solution5 Element, wt% C Si

Alloy

304L

18.39

1.55

10.5

0.026

0.47

0.004

0.08

304LN

18.87

1.58

9.4

0.025

0.35

0.004

0.16

Behind Van Stone ﬂanges (external) Near welds Water ingress and accumulation of chlorides In cold-bent elbows

On ﬂange faces beneath gaskets FIG. 16

Piping locations susceptible to SCC.2

FIG. 17

Diagram of the Van Stone flange geometry.4

HYDROCARBON PROCESSING MARCH 2012

I 43

SPECIALREPORT

CORROSION CONTROL

SS foil wrapping of the column should be completed before installing the insulation on the entire external surface of the column, including the stiffening ring. Insulation material selected for SS equipment should be acceptable as per ASTM C795. To resist the chloride attack, designers should address the insulation chemical constituents (Na + silica) content. Forward path. Other recommendations were identified in the review process: 1. Conduct pre-monsoon insulation audits to check for openclad joints. Wherever the joints are not water tight, sealing putty should be applied. 2. Insulation at the stiffening ring in other SS columns in similar process conditions should be removed before carrying out inspection activities. 3. To avoid water logging at the stiffening ring, an intermittent welding process that is acceptable per clause UG-30 of ASME Sec. VIII Div. 1 should be deployed while welding the stiffening ring with the column. 4. Suitable paint should be applied on the column shell externally, as an extra precautionary measure to mitigate corrosion under insulation. 5. SS foil should be completely wrapped over the column before installing the insulation. The SS foil works as an additional barrier to avoid rainwater carrying leachable chlorides to the column shell. 6. Review the requirement of refined mineral wool to insulate SS equipment. HP

ACKNOWLEDGMENTS The author sincerely thanks the Reliance management for its assistance in publishing this article. LITERATURE CITED Cottis, R. A., Guides to Good Practice in Corrosion Control, Corrosion and Protection Centre, UMIST, UK, 1982. 2 Hira, A., Stress corrosion cracking post, http://csidesigns.com/flowgeeks/ stress-corrosion-cracking. 3 “Chloride Stress Corrosion Cracking of Buried Stainless Steel Pipeline,” Corrosion Testing Laboratories, www.corrosionlab.com. 4 Definition and details of flanges, Explore the world of piping, www.wermac. com. 5 “Standard Practices for Detecting Susceptibility to Intergranular Attack in Austenitic Stainless Steels,” ASTM A262 Practice A. 6 “Standard Test Methods for Chemical Analysis of Thermal Insulation Materials for Leachable Chloride, Fluoride, Silicate, and Sodium Ions,” ASTM C871. 7 “Standard Specification for Thermal Insulation for Use in Contact with Austenitic Stainless Steel,” ASTM C 795. 8 “Rules for construction of pressure vessels,” ASME Sec VIII Div. 1. 9 Anita, T. and N. S. Bharasi, “Stress corrosion cracking behavior of AISI Type 304LN stainless steels with different nitrogen contents,” IGCAR, Kalpakkam, Tamil Nadu, 2010. 1

Anjul Deep Jain is the manager of the Engineering Division of Reliance Ports and Terminals Ltd., at Navi Mumbia, India. He began his career with Reliance Industries Ltd., in 2008. His areas of specialty include corrosion and inspections, nondestructive examinations and failure analysis of inservice equipment. Mr. Jain holds a B. Tech degree in metallurgy and material science engineering from the Indian Institute of Technology.

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CORROSION CONTROL

SPECIALREPORT

Operating philosophy can reduce overhead corrosion Boost refinery reliability by controlling potential amine recycle loops M. DION, B. PAYNE and D. GROTEWOLD, GE Water and Process Technologies, The Woodlands, Texas

alt fouling and associated corrosion in the crude unit overhead are complex phenomena that impact refinery reliability, flexibility and, ultimately, profitability. Establishing an appropriate balance of physical, mechanical and operational parameters, unique to each unit, is critical to minimizing fouling and corrosion throughout the crude unit. Factors such as amine chloride salt points; optimum accumulator pH; and overhead water ICP (initial condensation point, also referred to as water dew point) are interrelated and all affect the potential for system fouling and corrosion. Further improvements to refinery reliability can be attained by controlling potential amine recycle loops that can cycle up amine concentrations and move the salt point upstream or within the atmospheric tower itself.

Crude unit overhead corrosion control. The first line

of defense against overhead fouling and corrosion is the desalter. The desalter is designed to remove the majority of water extractable chlorides that contribute to the formation of highly corrosive hydrochloric acid (HCl) in atmospheric overhead systems. Depending on the desalter design and operation, it typically extracts 90–98% of the water extractable species. To protect the system from extractable chlorides that are not removed in the desalter and non-extractable, hydrolysable chlorides (such as organic amine chlorides), filmers, neutralizers and an overhead water wash are commonly utilized. The first area of concern for overhead corrosion protection is at the initial condensation point (ICP). As the first drop of water condenses (Fig. 1), acids in the vapor phase will transition to the water droplet, creating a low pH, highly corrosive liquid. The neutralizing amine (N) must be present at the ICP to neutralize the hydrochloric acid. Amines can also associate with chlorides in the vapor phase under certain partial pressures, creating amine chloride salt. Once formed, it can migrate from the vapor phase either as a liquid or a solid and is typically extremely corrosive. The temperature at which the amine chloride salt exits the vapor phase is commonly referred to as the “salt point.” The salt point is dependent upon several factors, including the partial pressure of the neutralizing amine, the partial pressure of HCl, and the partial pressure of “tramp amines.” Tramp amines are generally defined as those other than neutralizer amines. They can come from several sources including

being present in the crude naturally; from upstream additives such as corrosion inhibitors or hydrogen sulfide scavengers; from another processing unit; or from compounds that may decompose into amines in the crude unit furnace. Control. Overhead pH control is, perhaps, the most important

aspect of overhead corrosion control. The pH in the overhead receiver is generally at least 0.5–1.5 points higher than the pH at the ICP. The ICP should be maintained in a range between 5.5 and 6.5 by use of an appropriate neutralizing amine. As illustrated in Fig. 2, operating at a pH level outside this range can have a deleterious impact in both directions. For example, if the accumulator pH is 5.5, the ICP pH will typically be between 4 and 5. When the ICP pH is 4.5 or less, acidic corrosion becomes very aggressive. Conversely, when the ICP pH exceeds 6.5, a region exists where the deposition of liquid or solid amine chloride salts can increase the likelihood of salt fouling and under deposit corrosion. H2S and other weak acids will increase partitioning from the vapor to the liquid phase as the pH increases. The additional sulfides and weak acids in the condensed water will act as a buffer requiring significantly more amounts of neutralizer for minor movements in pH. The additional neutralizer concentration increases the partial pressure of the neutralizing amine, thereby increasing its salt point and the associated risk for under deposit corrosion. Acids and bases at dew point H CI

H CI Henry partitioning

H CI

H CI N H+

N CIH+

H CI

H+ CIH+ N

CI-

Electrolytic chemistry

First water drop at ICP FIG. 1

Water chemistry for the initial condensation point.

HYDROCARBON PROCESSING MARCH 2012

I 45

SPECIALREPORT

CORROSION CONTROL

Additionally, the destruction of metal passivating iron sulfide scales also becomes a factor under these conditions. In a slightly acidic environment, sulfides will react with the iron, forming a protective iron sulfide film. This protective film is weakened as pH increases, inhibiting the effectiveness of the naturally occurring protective iron sulfide film. Therefore, both the upper and lower levels should be considered hard limits not to be exceeded. Having a pH excursion beyond these limits is generally an indication that there is a significant imbalance in the system from either an incidental or a systematic situation. Most refiners employ an overhead water wash to force the condensation of water vapor and dilute the acids that condense with the water. However, this may not protect against amine chloride salt fouling if the amine salt forms above the overhead temperature at the water wash injection point. The potential corrosion risk can also be compounded if the high salting amines reenter the atmospheric column in the reflux, which can induce an amine recycle loop. Amine recycle. As discussed previously, amines can be pres-

ent as either tramp amines or introduced into the overhead as neutralizing amines. When exposed to a liquid-liquid system, amines—such as monoethanolamine (MEA), diethanolamine (DEA), methyl diethanolamine (MDEA) and ethylenediamine

Corrosion rate as a function of ICP pH, mpy

1,000 Salt deposition; under-deposit corrosion (NOTE: This saltpoint curve will shift with varying amine and chloride concentrations)

900 800

Iron sulﬁde scale weakens and H2S/CO2 partitioning to liquid phase is enhanced

700 pH at which saltpoint exceeds water dewpoint Optimal control range pH 5.5 – 6.5

600 500 400 300

(NOTE: The corrosion rate at pH >7 is equivalent to the rate at pH 4)

200 100 0 1

(EDA)—will partition to each phase. For instance, in the overhead accumulator, a portion of the MEA will partition to the naphtha reflux and another portion will partition to the condensed water. If the condensate is used as desalter wash water, it will again partition, with a portion of the amine exiting the desalter in the desalted crude. This creates amine recycle loops (Fig. 3) in the naphtha overhead and desalted crude. These recycle loops can concentrate the amine within the system. The additional amine loading to the overhead will add to the partial pressure of that particular amine, which will, in turn, increase the salt point of the amine chloride salt. If left unchecked, this amine recycle loop may, in severe cases, foul the top distillation trays. Amine partitioning. The partitioning of amines between the hydrocarbon and water phase is dependent on many factors including the type of amine, the hydrocarbon polarity and the pH of the water. Low pH water can protonate (add protons to) an amine and drive the ionic compound into the water phase. Conversely, alkaline water will deprotonate an amine and drive the partitioning of the non-ionic compound into the hydrocarbon phase. Amine partitioning is dependent on the type of amine (Fig. 4). As more carbons are added to an amine compound, its partitioning will be less pronounced with pH. Ammonia is easily partitioned to the water phase; MEA partitions to a lesser extent; and so on. In a crude unit overhead, operating the overhead accumulator water at a slightly acidic pH will assist in breaking the reflux amine salt recycle loop. The use of a low salting amine to control pH at the initial condensation point and not salt above the water dew point is critical to an effective overhead corrosion control program. At the desalter, reducing the effluent brine pH will also drive more amines into the effluent brine, thereby minimizing the potential harmful impact from amine recycle loops. It should be noted that the effluent brine pH is the equilibrium pH after the Acid

Ammonia:

NH4+

NH3 Base

The impact of pH at the initial condensation point.

Neutralizer Amine sources include: t 0WFSIFBE OFVUSBMJ[FST t $SVEF PJM t 4MPQ PJM t "MLBOPMBNJOF VOJU t 4PVS XBUFS TUSJQQFST t )24 TDBWFOHFST t $PME XFU SFnVY

Amine recycle

Fractionation column

Wash water

HO – CH2 – CH2 – NH2

Base

HO – CH2 – CH2 – NH3+

Amine partitioning is dependent on the type of amine.

HDS efﬂuent exchanger dP (indication of exchanger plugging) Unit shutdown for cleaning Desalted pH modiﬁcation treatment Untreated

Eff dP actual Eff dP model 60

Tank farm

FIG. 3

Accumulator

Tower top reﬂux

Amine

Desalter

FIG. 4

Water wash

Amine recycle

Acid MEA:

psi dP

FIG. 2

Typical amine recycle loop diagram.

I MARCH 2012 HydrocarbonProcessing.com

Stripping steam

Untreated baseline to 3/2008 20 1/7/2011

FIG. 5

2/26/2011

4/17/2011

6/6/2011

7/26/2011

9/14/2011

Detailed rendering of the diesel hydrotreater effluent exchanger pressure drop.

CORROSION CONTROL

Know-how for Your Success

crude oil and wash water mix. Consequently, the effluent brine pH is the control parameter to amine partitioning within the desalter. Out at the refinery. A US refiner was experiencing through-

put reductions and frequent slowdowns as a result of fouling in the effluent side of the diesel hydrotreater feed effluent exchangers. Rather than treat the symptom with an amine halide salt dispersant, the desalter effluent brine pH was lowered by injecting a product containing citric acid and a scale inhibitor. This partitioned more amines to the effluent brine, reducing the amines in the crude unit overhead, the diesel stream and, consequently, the fouling in the hydrotreater unit. The effluent exchanger pressure drop history was used to generate a multiple regression linear model to normalize the pressure drop for effluent flow and stream properties. The actual exchanger pressure drop and the model estimate are shown in Fig. 5. When the actual pressure drop increases above the model’s predicted value, it is due to the amine halide salt fouling at an advanced rate. The time periods in Fig. 5 (during treatment) show that the actual pressure drop was lower than the historical observations and, in fact, there was no increase in pressure drop. Wrapping it up. In some systems, amines may recycle back

into the tower with the reflux or may reenter the desalter from the overhead condensate. At the desalter, the amines may partition back into the desalted crude and reenter the atmospheric tower. These amine recycle loops may cycle up amine concentrations and increase the risk of corrosion from amine salt deposits if they occur above the water dew point. The authors believe a model can be used to assist in predicting amine salt points. If the salt point occurs above the water dew point, operating the desalter with an acidic effluent brine can partition a portion of these amines into the effluent brine, thereby reducing the detrimental impact from recycling amines. Using nonvolatile acid products is a good way to assist in reducing the desalter effluent brine pH. The acid decomposes to inert substances in the crude unit furnace. As refiners have recently reduced atmospheric tower top temperatures to maximize diesel production, a thorough understanding of the ICP, salt point and control of amine recycle loops is critical to maintaining plant reliability in changing plant operational conditions. HP

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Delbert R. Grotewold is a senior regional engineer for GE Water and Process Technologies in the Western US region. He is responsible for process chemical treatment programs, refinery process troubleshooting and process optimization of refinery operations. Mr. Grotewold has over 27 years of refinery engineering and process treatment experience. He has authored or co-authored five patents in refinery technology. Select 162 at www.HydrocarbonProcessing.com/RS 䉴

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CORROSION CONTROL

SPECIALREPORT

Closed-loop control can clamp down on crude unit corrosion Automating the detection process and controlling applications in real time dramatically improves performance N. P. HILTON, Nalco Energy Services, Sugar Land, Texas

Traditional corrosion control. Refiners assess the corrosive

conditions of the crude atmospheric overhead on an arbitrary but set frequency. This can differ greatly between refineries. Best-inclass refiners are typically assessing the corrosive environment once a day, typically on the day shift. This is a reasonable approach during periods of stable operation, but this approach does not account for changes to operation or periods of unstable operation. The two major flaws to this approach are the frequency of data collection and the frequency of corrective actions taken. The purpose for accessing the corrosive environment is to understand the impact that changes to operations are having on the health of the overhead, as well as, the corresponding demands these changes are placing on the chemical treatment program. It stands to reason that changes to operation and adjustments to the chemical treatment program are only made when actionable data is available. In this model, the occurrences of data collection are far too infrequent, given that there can be large swings in the corrosive conditions during times of unstable operation. This is what should be referred to as “the corrosion window.” Refiners who constantly process a variety of crudes and have frequent periods of unstable operation need a better way of assessing the crude overhead conditions and making adjustments to operations immediately, not waiting hours for the refinery’s central laboratory to process samples collected overnight so they can have the data needed to take a corrective action. New solutions are being deployed that combine the capabilities of detecting changes in corrosive conditions and automati-

cally adjusting chemical and/or caustic addition to meet system demands in real time. For example, a new crude unit analyzer has been deployed that continuously monitors the pH, chloride and iron in refinery overhead sour water. The analyzer takes these results and automatically controls the addition of the neutralizing amine, sodium hydroxide and filming amine to meet the system demands within the control limits specified. Testing frequency. Refiners and specialty chemical suppliers typically assess the corrosive conditions of the crude overhead by conducting a handful of wet chemistry tests. These should include pH, chloride, iron, ammonia and sulfide as a minimum. These tests have been around for many years and are widely adopted by industry for their practicality. The problem is not the validity of the tests conducted, but the frequency of the data collected. In the best cases, refiners are interrogating their overhead systems once in a 24-hour period. In the worst cases, refiners are interrogating their overhead systems once per week. The measurement of pH, in many cases, will be the only exception to this. As a general practice, pH is measured at least once per shift. The other tests—chloride, iron, ammonia and sulfide (all good indicators of the corrosive conditions)—can be more involved and typically require the use of the refinery central laboratory or involvement of the specialty chemical supplier. As a result, these tests are not conducted as frequently. The time lag between sampling and an actionable result can be many Analyzer beta test data

80 70 60 Chloride in ppm

he refining industry has been dealing with crude unit overhead corrosion since the early days of the industry. Many refiners today understand that 90% of corrosion damage occurs 10% of the time, when persistent and ongoing overhead corrosion is prevalent. These small periods of corrosion are related to unstable operations, the processing of opportunity crudes or other interruptions in normal operations. The industry has historically tried to control the corrosive environment through the application of best practices; by using specialty chemicals including sodium hydroxide (caustic); by improving desalter operations; and by upgraded metallurgy. The gap with this approach is not the tools themselves but the application of these tools. Today’s industry best practices impact 90% of stable operations but fail to address the 10% of time during unstable operations when the lion’s share of corrosion damage can occur.

50 40

Cl Fe pH 1 pH 2

30 20 10 0 9/7/09 9/8/09 9/8/09 9/8/09 9/8/09 9/8/09 9/9/09 9/9/09 9/9/09 Date and time

FIG. 1

The analyzer beta test data shows that chloride and iron increased significantly with only a subtle change in pH. HYDROCARBON PROCESSING MARCH 2012

I 49

SPECIALREPORT

CORROSION CONTROL

hours, depending on who is conducting the test. Often, the upset condition has passed well before a result and anticipated actions are ready. This may be adequate for those refiners who are running a limited or restricted crude diet and experience few changes or periods of unstable operation; but, for the majority of refiners today, this is not optimal. Most refiners in this day and age are constantly juggling opportunity crude cargoes and managing frequent changes in operation, which increases the risk and frequency of periods of unstable operation. These refiners need to be accessing the corrosive environment on a far more frequent basis if they are to gain an understanding of the impact these changes are having on their equipment and reliability. TABLE 1. Three phase technology implementation over a 30-month period Phase

Time period

Description

Base conditions

2006 to March 1, 2009

Manual data collection/ Manual control

Monitoring phase

March 2, 2009 to September 17, 2010

Crude unit analyzer/ Manual control

Control phase

September 18, 2010 to present

Crude unit analyzer/ Auto control

Neutralizer pump Chloride ppm pH average

8 7

80 70 60 50 40

20 4

3 40452 40452.5 FIG. 2

40453 40453.5 40454 40454.5 Axis title

TABLE 2. A comparison of Probe 1 and Probe 2 from 2006 to 2011 Chloride ppm, neutralizer pump %

Crude overhead analyzer control phase

Todayâ&#x20AC;&#x2122;s practices tend to facilitate the feeding of chemical additives (caustic, neutralizing amine and filming amine) to a base-line dosage or to stable operations 90% of the time. Under current practices, refiners are not capable of measuring and/or responding to system changes and unstable operations fast enough. The result is that the right amount of chemical is rarely injected at the right time. Hence, things are constantly in a perpetual state of over-feeding or under-feeding chemical additives. The goal should be to always have the correct amount of chemical injected at the most opportune time and in the right location. If periods of unstable operation are caught, it is unlikely that the catch happens immediately. What usually happens is that upset conditions are detected in the middle of the upset, when the conditions are highly visible with much of the corrosion damage in progress. The result is a corrective action that is after the fact, resulting in only partially addressing the problem. One result of addressing upset conditions after the fact is that there is a tendency to feed high amounts of chemical to correct the situation. This inefficiency results in higher volumes of chemical consumed, often with little impact on the final result. The other more serious concern is smaller, subtler changes to operations that are not highly visible but are quietly damaging the integrity of the plant. One such example is captured in Fig. 1. What can clearly be seen from this graph is that only chloride

Probe 2

8.98

10.57

2007

7.83

9.07

2008

5.56

5.01

2009

7.90

8.46

5.36

4.79

2010 Monitor Mode

6.55

5.23

2010 Control Mode

2.67

3.59

2011 Control Mode

3.35

2.77

0 40455 40455.5

Corrosion excursion control mode

Alpha test site metal loss

2.00

A sharp increase in chloride is evident with a corresponding decrease in pH.

Metal loss, mil Running linear corrosion rate Chloride Linear (metal loss, mil)

Average rate 2.8 mpy

1.90

1.85

1.80

1.75

1.70

Chloride ppm Corr. average mpy pH average Iron ppm

40 30

Metal loss, mil

4 3

FIG. 3

Corrosion increased from a chloride and pH excursion.

I MARCH 2012 HydrocarbonProcessing.com

FIG. 4

The same data from Fig. 3 is reconfigured, with the corrosion rate expressed as metal loss.

10/6/10

10/4/10

10/2/10

0 40452.38 40452.88 40453.38 40453.88 40454.38 40454.88

9/30/10

-3 9/28/10

1.60 9/26/10

9/24/10

9/22/10

0 9/20/10

1.65 9/18/10

9/16/10

24-hour CR, mpy

Chloride ppm, Corrosion rate, mil per year, mpy

Probe 1

2009 Monitor Mode

1.95

Year 2006

CORROSION CONTROL and iron increased significantly, with only a subtle change in pH. In this instance, if the only gauge of corrosive conditions was pH, operations and most likely the specialty chemical supplier would not have taken any corrective action. However, as the graph shows, there is clearly a corrosive event taking place here that is not represented by the pH of this highly buffered system. Solution. The key to controlling crude overhead corrosion

is twofold. The first goal is to collect accurate data frequently enough to be able to track unstable operating conditions from the onset through the duration of the event, while simultaneously assessing the severity. After that, the ideal arrangement would be to directly link the addition of chemical additives based on the systemâ&#x20AC;&#x2122;s true demands to these periods of unstable operations. For the past 30 months such an analyzer has been deployed in a North American refinery. Process sour water from the crude overhead has been continuously sampled for pH, chloride and iron. These results are then simultaneously stored in a process historian and run through a commercially available programmable logic controller (PLC). The function of the PLC is to assess the measured result and take action based on the results in real time. As with any modern sophisticated controller, the means to alarm and calculate many important parameters exists. The ability to control the addition of chemical additives to best practice

standards and then constantly adjust the dosage to meet system demands is real. It is now possible to directly link a change in desalter operations with performance in the overhead. In Fig. 2, a sharp increase in chloride is observed with a corresponding decrease in pH, while the analyzer is simultaneously adjusting the neutralizing amine dosage to meet the system demands. The function of alarming key overhead parameters (such as detecting large swings in pH and chloride) can now be fed directly back to the unit operators and the specialty chemical provider, alerting them to a potential onset of an upset condition. The process historian now provides the ability to go back in time to help develop a better understanding of process changes and their impact on the crude overhead system. The data is available in real time to unit operations, engineering and the specialty chemical supplier. As with any system, the true indication of performance is the corrosion rate. During the 30-month trial period, electrical resistance probes were located on the inlet to the first two overhead exchangers. These probes had historically been good indicators of the performance of the overhead. For the first 12 months (and prior to implementing full control of the addition of chemical additives), these probes were manually read once a week. The problem with this approach was that there was not enough clarity to detect daily unit changes. What was needed was the ability to

TABLE 3. Quantification of the corrosion rate reduction for both probes

8.98

10.57

2007

7.83

9.07

2008

5.56

5.01

2009

7.90

8.46

2009 Monitor Phase

4.17

4.35

2010 Monitor Phase

6.55

5.23

2010 Control Phase

2.67

3.59

2011 Control Phase

3.35

2.77

Monitor vs Control

Probe 1 Probe 2

10 Corrosion rates, mpy

Year

% Improvement % Improvement Probe 1 Probe 2 Probe 1 Probe 2 mpy mpy vs 2006 vs 2006

2006

SPECIALREPORT

8 6 4 2

40%

55% 0 2006

66%

70%

44%

34% FIG. 5

2007

2008

2009

2009 2010 2010 2011 Monitor Monitor Control Control mode mode mode mode

A comparison of Probe 1 to Probe 2 from 2006 to 2011.

TABLE 4. A summary of inspection records for Exchanger 1 and Exchanger 2 $25,000

Inspection history 1995 to 2011 Year Exchanger 1

Year

Exchanger 2

Dec-95

Bundle replaced

Jan-96

Bundle replaced

Oct-97

Retube C-E-1D installed

Oct-97

Inspection

Oct-98

Retube

Dec-99

10 leaking tube

Dec-99

514 tubes plugged

Aug-00

14 leaking tubes

Mar-00

Bundle replaced

Jan-02

120 tubes plugged

Oct-02

Passed inspection

Mar-03

235 tubes plugged

Jan-04

Retube

Jan-04

1 tube plugged

Apr-04

1 tube plugged

Apr-04

310 tubes plugged

Jul-06

Bundle replaced

Feb-05

Bundle replaced

Apr-07

Inspection no repairs

Jul-07

Inspection no repairs

Jun-07

Spare bundle installed

Mar-08

Spare bundle installed

Mar-08

1 tube, 60% wall loss

Monthly average $20,000 $15,000 $10,000 $5,000 0 2006

FIG. 6

2007

2008

2009

2009 2010 2010 2011 Monitor Monitor Control Control phase phase phase phase

Average monthly spending for both neutralizing amine and filmer since 2006. HYDROCARBON PROCESSING MARCH 2012

I 51

SPECIALREPORT

CORROSION CONTROL

detect unit upsets and corrosion episodes in real time. To facilitate this, the probe signals were brought directly into the analyzer, allowing for an instantaneous corrosion rate. This data was then pushed to the process historian and logged. Fig. 3 shows the resulting increase in corrosion from a chloride and pH excursion. Here the corrosion rate is expressed in terms of mil per year (mpy). Fig. 4 presents the same data in a different light. Here the corrosion rate is expressed as metal loss (mil). It is clear that the base corrosion rate is fairly low, but there is a sharp and clear change in slope corresponding with the two excursions. Both these graphs are evidence that corrosion damage is accelerated during times of unstable operation. Impact(s). To understand the impacts experienced by the refinery over the past 30 months, one needs to understand that there have been several phases to the implementation of this technology (Table 1). The results achieved show improved performance in all phases during the 30 months. Not only has the refiner seen performance improvements and increased the life of the overhead bundles, but these results were achieved with a reduction in chemical additives and caustic spend, a true win/win.

TABLE 5. Incoming and desalted crude content from 2006 to 2011 Salt in, ptb*

Salt out, ptb

2006

8.7

0.7

2007

12.2

0.8

2008

16.9

0.8

2009

21.1

2.1

2009 Monitor Phase

16.1

1.1

2010 Monitor Phase

12.1

0.9

2010 Control Phase

12.4

0.7

2011 Control Phase

18.4

*Note: ptb stands for â&#x20AC;&#x153;pounds per thousand barrelsâ&#x20AC;?

The refiner in question witnessed an overall improvement of 64% in corrosion on Probe 1 and 72% on Probe 2 when compared to base conditions. This was achieved with higher overall crude incoming salt levels and a spike in desalted crude salt levels during the 30 months. Fig. 5 and Table 2 break this data down further. During the initial monitoring phase, a 40% reduction in corrosion rate for Probe 1 and a 55% reduction in corrosion rate for Probe 2 were observed. During the control phase, Probe 1 saw an additional corrosion rate reduction of 44%, with Probe 2 reducing an additional 34% (Table 3). It is clear that the hourly performance data, coupled with extensive monitoring from onsite personnel, had a significant impact on the corrosion rates during the initial monitoring phase. What was uncertain is if performance would improve further during the control phase. The second step change was, without a doubt, due to the automatic control of the chemical additives during the control phase. The PLC was able to respond automatically to changes in pH, chloride and iron levels, thus delivering precisely the appropriate amount of chemical in the right place at the right time. As further validation, one only needs to look into the inspection records for these two exchangers. Table 4 summarizes the inspection records for both Exchanger 1 and Exchanger 2. The inspection records show that Exchanger 1 was new in 1995; retubed in October 1998; re-tubed again in January 2004; replaced in July 2006; and then swapped with a spare bundle in March of 2008. The longest period without downtime or maintenance was 24 months. Since March 2008, Exchanger 1 had been online for 44 months without maintenance or down time. Exchanger 2 had a similar story, although it saw less maintenance and downtime when compared to Exchanger 1, with 48 months being its longest period of sustained operation. Not only did automated control improve corrosion performance, but the performance was achieved with substantially lower costs for the chemical additives and caustic program. Fig. 6 represents the average monthly spends for both neutralizing amine and filmer since 2006. The chemical additives spend was reduced 42% from 2006 and 60% when compared to 2009

TABLE 6. A detailed accounting of expenses related to both exchanges from December 1995 to March 2008 Year

Exchanger 1

Year Jan-96

Exchanger 2

Bundle Labor and replacement/Retube inspection

Tube plugging

Lost production bbl

Total costs at $2/bbl upgrade

Dec-95

Bundle replaced

Oct-97

Retube Exchanger D installed Oct-97

Bundle replaced

$70,000

$350,000

320,000

$1,380,000

Inspection

$35,000

$175,000

320,000

Oct-98

Retube

$1,170,000

$35,000

$175,000

160,000

$690,000

Dec-99

10 tubes plugged

Dec-99

514 tubes plugged

Aug-00

14 tubes plugged

Mar-00

Bundle replaced

$75,000

$10,000

320,000

$1,045,000

$35,000

$175,000

$10,000

320,000

Jan-02

120 tubes plugged

Oct-02

Passed inspection

$1,180,000

$75,000

$10,000

320,000

Mar-03

235 tubes plugged

$1,045,000

$75,000

$10,000

160,000

$565,000

Jan-04

Retube

Jan-04

1 tube plugged

Apr-04

1 tube plugged

Apr-04

310 tubes plugged

$175,000

$10,000

320,000

$1,180,000

$75,000

$10,000

320,000

Jul-06

Bundle replaced

Feb-05

Bundle replaced

$70,000

$175,000

$1,045,000

320,000

$1,205,000

Apr-07

Inspection no repairs

160,000

$555,000

Jul-07

Inspection no repairs

Jun-07

Spare bundle installed

$35,000

$175,000

$10,000

320,000

$1,180,000

Mar-08

Spare bundle installed

Mar-08

I tube 60% wall loss

$35,000

$175,000

$10,000

320,000

$1,180,000

$350,000

$1,950,000

$80,000

3,680,000

$13,420,000

$35,000

$75,000

= Bundle replacement or Retube

**Source: Refiner inspection records** **assumption bundle retube $35,000/each, Labor for bundle replacement 5 x cost of bundle, if no bundle replacement cost $75,000, inspection and tube plugs $10,000, lost production 40,000/bpd 8 days x 2 exchanger 4 days x 1 exchanger, barrel upgrade $2.**

I MARCH 2012 HydrocarbonProcessing.com

CORROSION CONTROL prior to the installation of the crude overhead analyzer. A closer look at the increased spends for 2009 can be explained by the large increase seen that year in both the incoming and desalted crude salt content (Table 5). At each stage of the crude analyzer trial, improved performance was achieved with lower consumption of chemical additives and caustic (Fig. 6). Again, there is a clear step change from base conditions to the monitoring phase. The perfectly timed addition of the correct amount of chemical through the use of automation not only impacted performance but also allowed faster response to upset conditions, minimizing overdose and under dose conditions usually seen when operating under manual control. Total cost reduction. In order to assess the total cost of

operation, industry standard costs must be applied to the past performance of Exchanger 1 and Exchanger 2. Taking these costs and then multiplying by the failure and maintenance history has generated a total cost of $13.4 million since 1995. Thus, the refiner in question spent, on average, $1.03 million per year from 1995 to 2008 for equipment, labor and lost production. Table 6 summarizes the cost per year that the refiner incurred. Most refiners do not have the luxury of being able to maintain equipment on the run without paying some penalty in reduced unit throughput and the associated lost margin opportunity. Table 6 also summarizes the barrels of lost production during this period and factors these lost production barrels into the total costs at a $2/barrel upgrade margin. Granted, these are only the hard costs of equipment replace-

SPECIALREPORT

ment, labor and lost production. Other less tangible but very real costs associated with these failures—such as the increased safety exposure to employees and contractors and any potential environmental impacts of this work—are not captured in this analysis. Improved performance, but still a long way to go.

Refiners clearly still have a long way to go when it comes to automating and controlling many of the peripheral applications that can impact reliability. To compete in today’s highly competitive refining environment, refiners will need to deploy technologies such as the crude unit analyzer highlighted in this article if they want to capture every advantage possible. By any measurement, spending in excess of $1 million a year on crude unit overhead corrosion is an expensive proposition and one not easily undertaken given today’s current industry practices. What can clearly be seen from the performance data presented is that automating the corrosive condition detection process in the crude unit overhead and controlling the chemical additives and caustic application in real time have dramatically improved performance. The refiner in question has gone from spending in excess of $1 million a year to spending zero on this system since March 2008, all while lowering its spend on chemical additives and caustic by $174,000 a year. HP Nigel P. Hilton is the downstream marketing manager with Nalco Energy Services in Sugar Land, Texas. He joined Nalco in 1990, starting as a technical service representative. Mr. Hilton has held several positions throughout his Nalco career in sales and marketing in the US and Europe. His current responsibilities are the strategic development of new technologies for Nalco’s downstream refining and petrochemical division.

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CLEAN FUELS

BONUSREPORT

Choose a facility configuration based on financial metrics LP modeling can provide an unbiased, cost-based preview of a refinery/petrochemical plant design T. E. SWATY, Fluor, Sugar Land, Texas

he selection of a process configuration for a refining and petrochemical production facility is a task that requires both the skill of an engineer and the business acumen of a finance practitioner. The selection process usually requires facility yields and feedstock requirements, relevant prices, unit capacity required, utilities consumption, capital costs for the configuration (adjusted for location, timing and location tax policies), and a financial model. The financial metrics for each of the candidate configurations are calculated and used to rank the different opportunities. This article explores the use of linear program (LP) modeling techniques to select process technology configurations based on economic drivers (margin and capital), with the inclusion of a simplified financial model built into the LP. The financial model is used to estimate the key metrics and give the project developer an initial view of the viability of the configuration. Configuration selection problem.

The demand for additional product in local and selected export markets are the usual drivers for new process capacity located in an area that offers economic benefits. The advantages of a potential facility location can include feedstock availability, cost, location, byproduct markets, etc. The project developer’s problem is how to procure—with minimum capital investment—the lowest-cost feedstocks to produce the needed finished products, and then generate sufficient economic returns to justify the project.

tions. The goal is to select—on an unbiased basis—the configuration that will meet the economic returns needed at minimum cost. The traditional methodology to perform this analysis is: 1. Construct an LP or other model of the facility that incorporates all of the potential process technologies envisioned for the project. The model should include the following: • Product specifications for transportation fuels, as well as forecast changes to these specs as a result of environmental requirements • Prices for crudes, other feedstocks and product prices in constant dollars on a consistent basis • Outside utility purchase availabilities and costs • Known or expected market limitations • Process technology capacity limitations • The latest crude assays for potential crudes processed, with sufficient detail to support product specifications of transportation fuels and process yield drivers. 2. Establish a set of case studies to evaluate each of the process technology Crude assays

Property relationships

Process technology

Crude composition

Elements

Initial estimates

configurations selected by the design team, including: • Configurations based on previous configuration studies • Input from subject matter experts (SMEs) • Developer company’s requests. 3. Run the model for each of the cases to determine the yields, gross margin and individual unit capacity needed. • Model process yields to determine sufficient information to create an income statement for each case. • Process unit capacities are based on the economic drivers and technology licensor limitations provided for the study. 4. Use high-level capital cost estimation procedures—i.e., cost curves. • Estimate the inside battery limits (ISBL) of each process technology and other ancillary units. • Estimate the ISBL cost at the selected site(s) using a location factor. • Apply a factor to account for the offsites, owners’ costs and contingencies. 5. Use a financial model to determine the metrics for each case.

Stream properties out of tolerance

Matrix

Extract information

PIMS SI Within tolerance solution

Traditional problem solution. The

process configuration that will produce the desired finished products has many varia-

Compare properties

FIG. 1

Recursion in a process industry modeling system (PIMS).

HYDROCARBON PROCESSING MARCH 2012

I 55

BONUSREPORT

CLEAN FUELS

TABLE 1. Assumptions used in model creation Available crude oils Middle East Light at $60/bbl: 100,000-bpcd max.

Middle East Heavy at $52/bbl: 100,000-bpcd max.

Middle East Medium at $53/bbl: 100,000-bpcd max.

Transport fuels produced Liquefied petroleum gas at $540/ton (t) Regular unleaded gasoline at $68/bbl: 50,000-bbl max. Premium unleaded gasoline at $69/bbl: 9,100-bbl max. Jet fuel at $74/bbl: no limit Ultra-low-sulfur diesel (ULSD) at $75/bbl: no limit Petrochemicals produced Polypropylene (homopolymer and impact grades) at $1,375/t: 360 kilotons per year (kpy) for both Ethylene glycol at $990/t: 685 kpy Ethylene oxide at $1,420/t: 100 kpy Polyethylene [linear low-density polyethylene (LLDPE) and high-density polyethylene (HDPE)] at $1,336/t: 365 kpy for both Styrene at $1,100/t: 630 kpy Propylene oxide at $1,500/t: 300 kpy Key transport fuel specs Gasoline

ULSD

• • • •

• Sulfur: 8 ppm max. • Flash point: 135°F min. • Cetane index: 47 min.

Research octane number (RON): 93 regular, 97 premium min. Sulfur: 5 ppm max. Benzene: 0.95 liquid volume percent (LV%) max. Aromatics: 34 LV% max.

Jet fuel

Internal fuel oil

• Sulfur: 2,000 ppm max. • Freeze point: −52.6°F max. • Smoke: 19 mm min. with naphthalenes at 3 LV% max.

• Sulfur: 1.0 weight max. • Specific gravity: 1.1 max

Process units permitted Refining limits

Petrochemical unit limits

• • • • • • • • • • • • • • •

• • • • • • • • • • •

Crude unit: 460,000-bpcd max. Vacuum unit: 195,000-bpcd max. Naphtha hydrotreater (HT): free Continuous catalytic regeneration (CCR): 71,000-bpcd max. Kerosine HT: free Diesel HT: free Residual fluid catalytic cracker (RFCC): 100,000-bpcd max. FCC gasoline HT: free Hydrocracker (HC): 75,000-bpcd max. Delayed coker: free Residual desulfurizer (RDS): 58,000-bpcd max. Alkylation: free MTBE: free Hydrogen plant: free Sulfur plant: free

• Use the appropriate tax, depreciation method, debt interest rates, percent equity and accepted metric calculation methods for each case. • Metrics should include internal rate of return (IRR), net present value (NPV), and after-tax payback (ATP) period. 6. Rank various configurations by capital expenditure (CAPEX) with the financial metrics to select the final configuration. This methodology has been used in industry and has provided project developers with sufficient data to make an informed process configuration selection. 56

I MARCH 2012 HydrocarbonProcessing.com

Ethylene cracker: 1,200-kpy max. Aromatics extraction: free Aromatics [benzene, toluene, xylene (BTX)] fractionation: free Xylene isomerization: free Transalkylation: free Paraxylene recovery: free Polypropylene: free Polyethylene: free Styrene monomer propylene oxide (SMPO): free Ethylene glycol: free Ethylbenzene: free

However, this approach does have potential problems and requires many model runs to fully explore the configuration possibilities. The usual start to a configuration study is an examination of what the company did in its last analysis for similar products. This can maximize the use of previous work and potentially lower the project engineering costs. An often-overlooked, but important fact is that every configuration study is based on a unique set of feedstocks, market conditions and economic drivers that can lead to different solutions.

A better approach to configuration selection. The use of modern LP

modeling systems with the inclusion of capital costs can provide a more balanced and less biased view of the process technology configuration selection. The following sections present a method of adding the impact of capital on the configuration selection problem, which is implemented in a process industry modeling system (PIMS). This system permits the user to model key points like ISBL as a function of capacity, economy-of-scale exponents, different stream factors for each process unit, differ-

CLEAN FUELS

BONUSREPORT

TABLE 2. Feedstocks and salable products Price Feeds

$/t

Rate $/bbl

tpd

bpd

Kilotons per year, kpy

Millions of $USD ($MM)/y

Middle East Light

82,407

4,108

1,805

Middle East Heavy

100,000

5,176

1,898

Middle East Medium

87,642

4,470

1,695

Coal

332

121

Methanol

389

232

Lime

Natural gas

200

1,803

658

132

Total feed

14,649

5,575

Products

kpy

$MM/y

Liquid petroleum gas

369

198

Regular gasoline

39,686

1,676

978

Premium gasoline

9,086

387

229

Kerosine/jet fuel

–

ULS diesel

123,176

6,040

3,372

Petroleum coke Sulfur

538

1,010

47 118

779

284

PP homo

1,375

986

360

495

PP impact

1,375

986

360

495

Ethylene glycol

990

1,847

674

668

Diethylene glycol

840

153

Triethylene glycol

1,167

LLDPE

1,336

849

310

414

HDPE

1,336

1,000

365

488

Styrene

1,100

575

210

231

Propylene oxide

150

1,500

274

100

Propylene

900

207

Butadiene

975

622

227

221

Paraxylene

843

1,150

2,009

733

Toluene

750

–

Mixed xylenes

725

–

Ethylene oxide

1,420

274

100

142

–

3,089

1,127

Benzene

750

1,113

406

MTBE

650

542

Refinery fuel

Gross revenues Gross margin

ent location factors, different owner’s costs and contingency factors. A calculation of IRR, NPV and ATP is also included, using consistent pricing and capital costs (same dollar basis) to give the developer an initial view of the project cost and economic viability. Methodology. LP solution techniques include the concept of “recursion,” which is really a version of the successive substitution method of solution to ensure that the model is optimized and converged. The concept is that the model starts with an estimated crude and feedstock composi-

tion and uses it to determine the yields and properties of all model internal streams. The LP solver then determines the optimal solution based on this data. The solution can have different crude/feedstock compositions and rates. The model internal stream yields and properties are then recalculated by the PIMS and compared to the previous values. The model is converged if all of the internal stream property differences and stream allocation dispositions are within the user-specified tolerances. This technique, depicted in Fig. 1, is used by all commercial modeling systems.

305

198

128

14,061

9,510

–

3,933

The capital costs are modified during the recursion cycle via the PIMS simulator interface, and are structured to impact the economic solution (objective function) to account for the differences in investment costs based on the process unit capacities determined in the last solution cycle. There is a major data difference in the LP that must be addressed: the capacity per calendar day in the LP vs. the capacity per stream day used in the capital cost estimation for the ISBL. The technique converts the capacities from barrels per calendar day (bpcd) to barrels per stream day (bpsd) HYDROCARBON PROCESSING MARCH 2012

I 57

BONUSREPORT

CLEAN FUELS

before estimating capital cost, and then back to bpcd before insertion into the LP model. The total investment cost (TIC) for

a given unit is then updated based on the new capacity calculated by the model. This is shown in algebraic form:

TABLE 3. Selected process unit capacities Capacity utilization Crude distillation unit (CDU) No. 1

Capacity selected Thousand tpcd Thousand bpcd 37.58

CDU No. 1 Vacuum unit No. 1

Max. capacity Thousand tpcd Thousand bpcd –

270.06 20.27

460 –

Vacuum unit No. 1

130.61

Naphtha HT

27.94

300

Aromatics reformer

26.32

71.50

Kerosine HT

–

200

Distillate HT

77.74

200

–

100

53.07

100

FCC Residual FCC Carbon burnt

0.75

195

Catalytic naphtha HT

–

300

Mild HC

Global HC limit

Delayed coker

–

100

Vacuum residue DS

Saturated gas plant

19.53

–

C4 isomerization

3.46

–

Sulfuric acid alkylation Sulfuric acid regeneration H2 plant

11.61

–

0.09

–

0.11 MMscfd

300 MMscfd

Amine unit

0.78

Sulfur plant

0.78

Tail gas treater

0.02

BTX extractor

2.93

100

BTX benzene tower

4.71

100

BTX toluene tower

3.19

100

BTX xylene tower

10.13

100

Toluene disproportionation

2.74

3.50

Xylene isomerization

7.18

200

Paraxylene recovery

9.21

Pygas HT

100

Pygas HT

2.47

Reformate splitter

2.47

100

Steam cracker

3.26

3.46

Butadiene extractor

1.15

–

Fresh xylene

2.41

–

Pygas splitter Bimodel polyethylene

–

19.15

–

0.58

1.73

Polypropylene

1.97

Ethylene glycol

1.74

1.92

Ethylene oxide

–

0.27

0.85

0.27

0.82

12.68

SMPO

LLDPE Propylene oxide Investment, billions of $USD 58

I MARCH 2012 HydrocarbonProcessing.com

PIMS model with cost estimates.

100 19.36

Unit TIC = [new capacity (bpsd)/ base capacity (bpsd)]x ⫻ base-capacity ISBL ⫻ factor Where x is the economy-of-scale exponent, ISBL is the inside-battery-limit cost of the process, and factor adjusts the costs for location, offsites, owner’s costs and contingencies. The sum of these costs for all process units in the configuration gives an estimate of the facility TIC. The TIC recovery for each unit is then the TIC calculated as above, divided by all of the following: the capacity in bpsd, multiplied by the stream factor, multiplied by 365 calendar days, multiplied by five (years). These values are captured using a PIMS utility row, which debits the economics (objective function) of the model. The methodology presented in this article uses a utility to capture the amount of cash needed on a daily basis to pay for the total installed process unit cost over a five-year period. This technique also uses PIMS utilities to report the IRR, NPV and ATP for the adjusted capital costs and process yields during the recursion step of the solution. This is the simplest method of transmitting solution information to the existing PIMS reporting structure. This technique also permits economics to drive the selection of the process technologies used in the configuration. The inclusion of the capital costs inside the LP greatly reduces the number of LP runs and provides a reasonable estimate of the IRR and other financial metrics for the selected configuration. The general configuration is designed to produce transportation fuels, monomers, aromatics and polymers, and has many processing paths open; few unit capacities are limited or at a minimum. The assumptions presented in Table 1 were used in the creation of this model. The units have to “pay” for the capital cost and operating costs when the capacity limitations are “free.” The units are also limited by the maximum product demand and licensor constraints. Other financial model assumptions used in this example include the following: • Max. investment: $13 billion (B) • Min. IRR: 13% • Discount factor for NPV: 10% • Depreciation: 14 years with 10% salvage value • Construction period: four years (yearly spending pattern of total installed cost: 24.4%, 43.2%, 25.2%, 7.2%)

CLEAN FUELS • • • •

Tax rate: 25% Offsites: 60% of ISBL Owner’s cost: 10% of ISBL Contingency factor: 20% of ISBL

Full-range naphtha

Pygas HTU

Splt

Pygas

Process Technologies. The capacity-

limiting process unit in the facility is the RDS, as it limits the crude rate and sets the RFCC capacity. The RFCC gasoline treater is not needed because the 77-ppm gasoline it produces can be blended with the other gasoline components to meet the sulfur specification of 45 ppm max. The RDS is more valuable in this case than the delayed coker because the RDS configuration produces more net liquid that can be upgraded to salable product, and the availability of relatively inexpensive natural gas and coal mitigate the value of petroleum coke as fuel. The process unit capacities required to produce the finished petrochemicals and transportation fuels are shown in Table 3. Financial Metrics. The financial metrics

for the LP model are compared below to those calculated by a stand-alone financial model with the same capital, yields, capital spending pattern, depreciation and other financial assumptions: 1. PIMS approximate method • IRR: 13.3%

Rafﬁnate

Kerosine Kerosine HDS 0

LLDPE HDPE

Polyethylene Homo impact plant EO/EG EO/EG Di/Tri PO SMPO Styrene Jet ULSD

DHT 78 Naphtha

CDU 270

Benzene Paraxylene

Aromatics plant 800 kpy Px

VDU 131

Hydrocracker 75

VGO

C2+ Light hydrocracker naphtha Unconverted oil

Flows are in kpy SWS No fuel import Capacity in thousand bpcd Amine treatment plant

FIG. 2

VRDS 58

Fuel gas Hydrogen plant 110 MMscfd H2S Sulfur plant 780 tpcd

C4 Isom

Light distillates

Diesel

Vacuum bottoms

Refinery yield comments. No jet fuel or petroleum coke was produced. The pricing supplied made it more profitable to produce only ULSD while still meeting the flash point requirement. The only aromatics produced were paraxylene and benzene. All of the intermediate aromatics were converted or blended into gasoline. The selected configuration for the refinery and petrochemical facilities is shown in Fig. 2.

Polyethylene plant

Steam cracker 1,200 kpy

Selected process configuration.

The results of the optimal process technology arrangement determined by the PIMS model using the economic drivers specified, the financial assumptions noted, and the technologies available will be presented in three sections: • Finished product yields, feedstock consumption and gross margin • Process technologies selected and their capacities • Financial metrics from the LP and from a detailed stand-alone financial model. The facility production of salable products and the feedstocks required to produce them are shown in Table 2.

BONUSREPORT

MTBE

MTBE Alkylation Alkylate 12

RFCC 53 LCO slurry

H2 T G T U

Sulfur

Selected configuration for refinery and petrochemical facilities.

• NPV at 10%: $6.5 B • After-tax payback: 5.1 years 2. Detailed financial model • IRR: 13.1% • NPV at 10%: $6.3 B • After-tax payback: 5.2 years 3. Accuracy delta • IRR: 1.5% • NPV at 10%: 3.2% • After-tax payback: 1.9% The financial metrics are not the same because of the following factors: • PIMS results contain non-sold items (nitrogen, spent lime, combustion gases) to material-balance the model • The detailed financial model excludes the non-sold items and only uses what is sold • These extra tons sold at a low price still impact the PIMS view of the configuration profitability and make the PIMS results higher for the IRR and NPV, but lower for the payback time. Takeaway. The inclusion of capital

costs and financial metrics in an LP model of a potential configuration provides the user with an unbiased view of the facility, based on economics. The identified

configuration is a viable starting point for the selection analysis and should be part of every feasibility study where the LP is a key analysis tool. This technique does not eliminate the need for sound engineering judgment nor the role of the SMEs. These are needed when the “optimal configuration” is converted to a practical solution with all of the elements not considered in the LP model. These elements include sour water stripping; practical unit capacities; multiple pressure steam levels; ancillary utilities; flares; and health, safety and environment concerns. HP

Timothy E. Swaty has more than 37 years of professional experience in most aspects of petroleum refining, and broad exposure to the petrochemical industry. He is currently a principal technical specialist with Fluor. His work at Fluor includes the application of management science techniques to refining and petrochemical problems, process optimization and financial analysis. Mr. Swaty has a BSChe degree from the University of Houston and an MBA from Texas A&M University (Corpus Christi). He is a registered professional engineer in Texas, Kansas and California. HYDROCARBON PROCESSING MARCH 2012

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Select 61 at www.HydrocarbonProcessing.com/RS

CLEAN FUELS

BONUSREPORT

Optimize hydrogen management for distillate production New tools enable refiners to fine-tune refinery configuration and maximize profits P. PARIHAR, R. KUMAR, R. K. VOOLAPALLI and S. AGARWAL, Bharat Petroleum Corp., Ltd., India

oving more distillate streams to diesel production and/or fuel oil (FO) production is a major refining activity. The distillate qualities and quantities, hydrogen price and consumption, product demand and prices, and refinery configuration constraints are key factors for these decisions. Producing higher-quality products from poorer-quality crudes requires more hydroprocessing. 1â&#x20AC;&#x201C;3 These distillate streams, e.g., heavy kerosine (HK), gasoils (GOs) and light cycle oils (LCOs), etc., are also used as cutter stocks to upgrade vacuum residues (VRs) and downgraded as low-value FO products. FO and VRs are major constrains if no resid-upgrading facilities are available. The heavy distillate streams from the fluid catalytic cracking unit (FCCU) must be hydrotreated to upgrade to diesel specifications. Hydrotreating consumes substantial hydrogen quantities, and this process adds more costs. When adding new hydrogenconsumer streams, hydrogen demand can exceed available refinery supplies.4 In such scenarios, optimizing existing hydrogen supplies is the key to improving the total profitability of the refinery. A simple reliable optimization tool for routing intermediate distillate streams maximizes benefits while meeting all constraints of hydrotreating capacity, hydrogen, VR utilization, product specifications and prices. The Excel-based calculator is generic for hydrogen management, and it also fits into various refinery configurations where different resid-evacuation options are practiced.

Hydrogen management. Processing opportunity crude oils and meeting critical product specifications of EURO III, IV and V are the real challenges for refiners. The quality of straight-run (SR) distillate streams obtained from the processing of high sulfur (S) and low API crude oils is considered inferior. Thus, refining higher-quality products from poorer-quality crudes has increased hydrogen addition. Conversely, LCOs (from the FCCU) are such intermediate distillate streams; they are also routed through the diesel hydrodesulfurization (DHDS) unit to meet product specifications. These streams (especially LCO) consume substantial hydrogen quantities, thus increasing processing costs to meet final diesel specifications. Under these conditions, optimizing hydrogen consumption is the key to total profitability. Planning tool. To fully optimize the routing of intermediate distillate streams and hydrogen management, additional requirements for these streams must be added to existing planning and

optimization tools. These tools must not only determine hydrogen and hydrocarbon routings, but also accommodate individual unit capacities and refinery configuration constraints. These intermediate distillate streams, in general, are also routed to residue evacuation when no modern resid-upgrading facilities such as coker/visbreaker/solvent deasphalting (SDA) are available. Adding intermediate distillate streams as cutter stock with residues produces FO, when capital investment may not be required.5 In the presented study, an optimization tool evaluates the routing of intermediate distillate streams to the diesel pool and/or resid-upgrading while meeting the refinery configuration constraints. More hydrogen demand. Hydrogen consumption for intermediate distillate streams has grown significantly.6â&#x20AC;&#x201C;8 Hydrogenaddition processes, in general, are preferred due to two factors. First, new environmental regulations over transportation fuels require higher-quality refinery products.9 Second, the differential prices for light- and heavy-crude oils continue to increase as lightcrude reserves are declining, and supplies of heavy-crude oils are increasing. Refiners are taking advantage of these spreads; they are incorporating more lower-cost, heavier, sour, opportunity crudes into the feedslate.10 Under these conditions, it is essential to understand the crude oils and their hydrogen content. Increased hydrogen consumption is an additional cost to process these crude oils. Therefore, to produce the same yields of transportation fuels either carbon rejection and/or hydrogen-addition processes must be selected. In actuality, even with incremental new carbon-rejection process capacity, additional hydrogen consumptions and, thus, their enhanced process capacities are preferred.11 This processing scheme enables optimizing hydrogen management for the refinery. The presented optimization tool can help facilitate efficient hydrogen usage in various resid-upgrading scenarios. Problem definition. While processing high-sulfur crude oils at a crude distillation unit, the processed distillate qualities were considered inferior. These intermediate distillates and LCO (from the FCCU) streams are routed through the DHDS unit to meet diesel-product specifications. These streams (especially, LCO) consume additional hydrogen to meet final diesel specifications. However, there are capacity limitations for hydrogen (maximum of 35 tpd) and hydraulic limits for DHDS capacity (maximum HYDROCARBON PROCESSING MARCH 2012

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BONUSREPORT

CLEAN FUELS

of 6,000 tpd). This is a common problem for any refinery. The presented study can be replicated to any other refineries with many commonalities and/or additional constraints. In the presented case, the VR is being evacuated as FO, where it consumes distillates as cutter stock to meet the final product specifications. Two grades of FO (180 cst and 380 cst) are produced when the cutter profiles are different.5 Thus, the available distillate streams are being used either for diesel production, which has a higher value and/or routed to FO production, which is needed for upgrading VR. The minimum VR production is approximately 3,500 tpd while processing 18,000 tpd of crude oils (6 MM tpy crude oil processing basis). The distillate streams available at this refinery are HK (high sulfur), GO (high sulfur) and LCO (high sulfur), etc. Fig. 1 shows the processing flow diagram for routing distillate streams. The cutter requirement depends on the quality of the produced distillate products. However, while optimizing the overall FO production, one must consider the total hydrogen consumption for upgrading both SR and cracked feedstocks. A simple optimization tool was developed for optimal routing of the distillate streams to diesel and/or FO production. The study considered all constraints of meeting DHDS capacity, hydrogen capacity, VR utilization and product specifications, prices, etc. METHODOLOGY

Diesel production is always the first choice due to its higher value over FO production. When distillate streams are routed DHDS

HK-diesel

Diesel-pool

GO-diesel HS-HK

HS-GO

LCO

LCO-diesel HK-FO

GO-FO

FIG. 1

Flow diagram of Base Case refinery.

FIG. 2

Screen shot of stream routing optimizer.

I MARCH 2012 HydrocarbonProcessing.com

Hydrogen consumption. In diesel hydrotreating, hydrogen consumption is governed by feed properties and product specifications. The affecting variables are carbon/hydrogen (C/H) ratio, S, basic nitrogen (N) and metal content, etc. To estimate the C/H ratio, a correlation as a function of specific gravity is applied.12 In this study, C, H and impurities (I) are evaluated in a balanced approach, especially across the DHDS unit to determine H2 consumption while upgrading distillates. The estimated H consumption is based on these assumptions: • Data for distillates are from refinery test runs • H2 consumptions are based on the distillate quality • All components, other than C and H, are considered as impurities for the calculations • Estimated cost for H2 is $2,150/ton. In this approach, the Excel-based optimization tool was developed to maximize benefits by optimum routing of intermediate streams to diesel (via DHDS) and/or FO production. The spreadsheet is enabled with macro, where input and output are linked with a single click button, as shown in Fig. 2. The model has provision to enter all inputs for the total distillate quantities available to routing and qualities, product specifications and prices, cutter profiles to meet product specifications and all process limits, e.g., DHDS capacity, hydrogen and VR upgrading, etc. These input data are treated as the Base Case, which is normally being practiced. The output data are reported as H2-consumption profiles for each stream, capacity utilization, optimum routing of intermediate distillates, final product profiles and overall benefits.

FO-LSHS

LCO-FO

through the DHDS unit, hydrogen is consumed to meet diesel-product specifications and, of course, increases processing costs. However, FO production needs no additional cost, but FO demand is declining. In this scenario, VR upgrading is one of the limits when equivalent distillates are downgraded. Thus, an optimal decision must be made between additional hydrogen consumption vs. producing more low-value FO products.

FO-LSHS

MODELING AND SIMULATION

Under present refining trends of processing various qualities of crude oils, distillates of varying qualities are produced by the CDU and secondary processing cracker units. Depending on product demand and prices, adjusting the processing scheme can increase profitability while complying with refinery limits. To consider global optimization with constraints, the presented model equations are generic to accommodate changes on a regular basis, and the model can be applied to other refinery configurations as well. The distillate streams are routed to DHDS, FO and excess. The excess term was introduced into the model equations for the mass balance. The intermediate distillate streams considered in the study are HK, GO and LCO. The possibility to evaluate effects from additional distillates obtained from different units and their routings can also be included in this model. Mass-balance equations for intermediate distillate streams are formulated based on their routings as:

CLEAN FUELS HK = HKDHDS + HKFO + HKExcess LCO = LCODHDS + LCOFO + LCOExcess GO = GODHDS + GOFO + GOExcess Mass balance around DHDS:

(1) (2) (3)

HDScapacity = HKDHDS + GODHDS + LCODHDS

(4)

Total FO production: FO = 100 HKFO /(cutter%)HK+100 LCOFO /(cutter%) LCO + 100 GOFO /(cutter%)GO (5)

BONUSREPORT

Optimization strategies. The key to any refinery profit

improvement program includes a single objective to maximize production while minimizing operating costs. The presented optimization program focuses on the present refinery practice of distillate routing for diesel and/or FO productions with the prevailing limits. The DHDS unit capacity, hydrogen capacity, VR upgrading and individual distillate stream capacities into the product pool are constraints considered in the study, as shown in Eqs. 11–16. 0.8700

VR mass balance equation: VR = (FO – HKFO – LCOFO–GOFO ) – (other cutters)

(6) (7)

H2DHDS = f [qualities (feed and product)] (8) H2DHDS (wt% of feed) =13.5 (DensityFeed – DensityProduct) + 0.31 Sulfurfeed – 0.20 (9) The listed equations are generic; new distillate streams can be introduced as variables to the model. Thus, the study can be extended to other various refinery configurations to optimize routing of intermediate distillate streams.

0.8600 Density, gm/cc

Mass balance for H2 consumption around DHDS: H2DHDS = H2GO + H2HK + H2LCO

0.1 wt% of feed 0.25 wt% of feed 0.5 wt% of feed 0.75 wt% of feed 1 wt% of feed

0.8650 Constant H2 consumption lines

0.8550 0.8500 0.8450 0.8400 0.8350 0.8300 0

FIG. 3

0.5

1.0

1.5 Sulfur, wt%

2.0

2.5

3.0

H2 consumption (wt% of feed to the hydrotreater) as a function of DHDS feed properties.

TABLE 1. Regression statistics for predicting H2 consumption (Eq. 9) Multiple R

0.998

R Square

0.996

Adjusted R Square

0.996

Standard Error

0.010

Observations

17 75 8

188

TABLE 2. Test run data of intermediate distillate streams HK

LCO

0.7992

0.8426

0.9061

0.33

1.25

0.28

> 105

Aniline point, °C

62.9

77.5

28.4

Pour point, °C

–40

IBP

179

256

166

5% vol Rec. @ °C

193

309

190

10% vol Rec. @ °C

193

321

199

20% vol Rec. @ °C

197

334

214

30% vol Rec. @ °C

201

343

228

40% vol Rec. @ °C

206

350

242

50% vol Rec. @ °C

210

356

257

60% vol Rec. @ °C

217

362

274

70% vol Rec. @ °C

224

367

292

80% vol Rec. @ °C

233

–

314

90% vol Rec. @ °C

244

–

341

95% vol Rec. @ °C

253

–

356

FBP

268

–

365

Density, gg/cc @ 15°C Sulfur, wt% Flash point, °C

Distillation

Base FIG. 4

Case 1

Case 2

Case 3

Case 4

Effect of varying feed composition on H2 consumption (%H2 unit capacity) during DHDS.

100 96 92 88 84 80 76 72 68 64 60

H2 consumption % of H2 plant capacity

Properties

100

H2 plant capacity utilization, % HK, w% GO, wt % LCO, wt%

HK LCO

0 2 4 6 8 10 12 14 16 18 20 22 24 26 28 30 32 34 36 Routing of intermediate distillates, HK & LCO towards DHDS, wt% LCO 17 wt% & HK 8 wt% LCO 12.5 wt% & HK 12.5 wt% LCO 0 wt% & HK 25 wt% FIG. 5

H2 consumption 100% of capacity H2 consumption 91% of capacity H2 consumption 68% of capacity

Effect of varying HK and LCO contribution in DHDS feed (unit capacity constant). HYDROCARBON PROCESSING MARCH 2012

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BONUSREPORT

CLEAN FUELS

Objective function: Maximize [(DHDScapacity DIESELYIELD% DIESELPRICE ) + (FO FOPRICE ) – (H2DHDS H2PRICE )] Constraints DHDScapacity ≤ 6,000 tpd VR ≤ 3,500 tpd H2DHDS ≤ 35 tpd HKDHDS ≤ 2,000 tpd LCODHDS ≤ 1,050 tpd GODHDS ≤ 4,500 tpd

(10) (11) (12) (13) (14) (15) (16)

Increase in FO production, wt%

Results. Table 2 summarizes the physical/chemical properties of the distillate streams from the CDUs (HK and GO) and excracker (LCO). These streams are typically used to meet demand for distillate production and/or used as cutter for VR upgrading. To meet product specifications, these intermediate streams are hydrotreated. The qualities of distillates for density, sulfur, flash point, pour point, etc., are critical factors in routing product streams. Thus, the qualities and quantities of distillates are important when optimizing the product routing to obtain maximum value for produced end products on a daily basis. In particular, while processing various crude oils, the availability of distillate quantities and qualities do change on a regular basis. However, the configuration and the capacity of the refinery are constant. In such scenarios, the feed to the hydrotreater can be optimized through the blending of distillates before processing 2.0 1.8 1.6 1.4 1.2 1.0 0.8 0.6 0.4 0.2 0.0

HK LCO

them. The variation in H2 consumption with throughput as feed (wt%) was studied as a function of feed density and sulfur. Inferior qualities of distillates have a higher density and higher sulfur content, and, accordingly, they consume more H2, as reported in Fig. 3. In the absence of surplus H2 capacity, blending distillate streams can be optimized to meet requirements for transportation fuel-product specifications. Among the three distillates, increasing HK proportion and reducing LCO proportion result in improved-quality feedstocks for the hydrotreater. In this approach, the combined feed to the hydrotreater has a lower density and lower sulfur content and it consumes less H2. As a result, H2 consumption (expressed as % of H2 plant capacity) is reduced, as shown in Fig. 4. The basis for this study is H2 consumption, which is equivalent to plant capacity (100%), where LCO (17 wt%) and HK (8 wt%) and GO (75 wt%) is the Base Case. To reduce the H2 plant capacity utilization, LCO proportion was reduced and HK proportion increased, while GO is constant (Cases 2 and 3). To further reduce H2 consumption, the GO proportion was reduced; this action yielded H2 savings up to 35% from the Base Case. Fig. 5 shows the detailed profile of routing HK and LCO to the hydrotreater. This figure clearly indicates the effects of simply manipulating qualities of feed routed to the hydrotreater. In the case of multiple feeds, optimizing feed qualities with H2 capacity is profitable. The qualities of distillates are critical factors for routing decisions. Thus, LCO is an inappropriate option for hydroprocessing, as compared to HK, since it consumes substantially more H2. However, when an LCO stream is routed to FO production, the entire LCO cannot be upgraded due to differences in the qualities of LCO and HK to meet the FO product specifications. However, to upgrade some of the VR, FO production can be increased, as shown in Fig. 6. Routing LCO to FO production may seem to be the obvious choice for FO production and H2 savings. Therefore, the re-routing of HK and LCO has reduced H2 consumption up to 32% and increased FO production up to 2%, and it complies TABLE 3. Optimization results for routing of intermediate distillate streams Component

0 FIG. 6

2 4 6 8 10 12 14 16 18 20 22 24 Routing of intermediate distillates HK&LCO, wt % towards FO Effect of LCO as a preferred cutter stock for FO production (Base Case-LCO 0 wt%, HK 23 wt%).

450

1,706

GO, tpd

4,500

4,294

LCO, tpd

H2 consumption, % of H2 plant capacity

Diesel, tpd

1,050

0.8505

0.8305

5,910

Routing to FO production, tpd

HK, tpd

1,100

GO, tpd

206

LCO, tpd

1,050

VR evacuated, tpd

3,500

FO, tpd

4,867

5,020

HK, tpd

2,450

2,294

GO, tpd

LCO, tpd

–

70 65 0.0

FIG. 7

Hydrogen, tpd

Optimum

HK, tpd

Combined density, g/cc 100

Base

Routing to DHDS, tpd

Excess, tpd 0.5

1.0 1.5 Increase in FO production, %

2.0

Effect of routing LCO (for VR upgrading) on H2 consumption and FO production.

I MARCH 2012 HydrocarbonProcessing.com

2.5

Additional revenue, $ million/yr

CLEAN FUELS

CASE STUDY

Scheduling and blending activities continue to face a dynamic landscape toward optimization and thus, refinery profitability. These variances include feedstock availability and qualities, product demand, resource limitations, configuration constrains and environmental regulations for fuels. Under the present trends of processing difficult crude oils to get the initial discount benefits and to meet tighter demand of fuel specifications, it is truly an optimization challenge. Modeling the routing of distillate streams was based on current practices at refineries. Model Eqs. 1–16 include distillate quantities and qualities, H2 consumption and capacity utilizations. The model equations are generic and can be tailored site-specific configurations. Based on modeling, simulation and optimization, an Excel-based tool was developed where all inputs and the Base Case can be entered. Optimal solutions can be compared with base values, and thus, the differential will be estimated as the benefit. A similar example is illustrated in Table 3. The optimization tool suggests that instead of routing LCO to DHDS, it is more beneficial to divert it to the FO pool. This is due to the increase in FO production and reduction in H2 consumption during hydrotreating. Accordingly, the HK and GO streams were routed to DHDS for upgrading to meet diesel specifications. In the H2-deficient scenario, diversion of streams can further be optimized by putting the additional constraints for H2 capacity while meeting the other important requirements. In turn, the surplus H2 capacity can be used to augment hydroprocessing capacity, and, thus, better H2 management is possible. This tool can be further developed as a custom-built refinery-wide scheduling and optimization utility. Options. Routing of distillate streams to diesel production is always the best choice. However, meeting all limits of hydrotreating capacity, H2, VR upgrading, availability of distillates and qualities, and product demand are essential. The excel-based optimization tool can enable H2 management while improving resources and capacity utilization more proficiently. In processing of opportunity crude oils and carbon rejection processes, the current optimization tool could be supportive for the optimum routing of intermediate distillate streams for maximum value. HP ACKNOWLEDGMENT The authors express their sincere thanks to K. V. Seshadri, ED (MR/R&D) of BPCL for his continuous mentoring on research activities and encouragement and many thanks to CRDC colleagues for constant appreciation and support. LITERATURE CITED Meenakshisundaram, A., “Hydroprocessing for fuels and lubes production,” Hydroprocessing, Bulletin of the Catalysis Society of India 3, pp. 1–9, 2004. 2 www.uop.com/.../UOP-hydroprocessing-innovations-supplement-techpaper.pdf. 3 Kumar, R., P. Parihar and R. K. Voolapalli, “New crude oil basket for hydrogen savings in refining,” Petroleum Technology Quarterly, Spring 2012. 4 Reynolds, B. E., J. L. Rogers, S. Spieler, and R. A. Broussard, “Resid Conversion Options,” No. 073, 1998, Chevron Research and Technology, Richmond, California. 5 Kumar, R., V. Chithra, V. C. Peddy V. Rao and N. V. Choudary, “Diverting 1

100

H2 consumption, % of H2 plant capacity

with the configuration limits, as shown in Fig. 7. The combined effect contributed up to 1% additional revenue, as shown in Fig. 8. In such a scenario, the availability of intermediate streams and their optimum routing are important when meeting the configuration limits on a regular basis.

BONUSREPORT

95 90 85 80 75 70 65 60 0.0

FIG. 8

0.2

0.4 0.6 Increase in revenue, %

0.8

1.0

Effect of intermediate distillate routing on H2 consumption and revenue.

low-sulphur heavy stocks from fuel oil production,” Petroleum Technology Quarterly, Summer, pp. 43–47, 2011. 6 Krenzke, L. D., J. E. Kennedy, K. Baron and M. Skripek, “Hydrotreating technology improvements for low-emissions fuels,” NPRA Annual Meeting, March 17-19, 1996. 7 Stanger, C. W., R. Fletcher, C. Johnson and T. Reid, “New process technology already existing in your refinery: Hydroprocessing-FCC synergy,” NPRA Annual Meeting, March 17–19, 1996. 8 Rhodes, A. K., “Distillation Capacity Exceeds 76 Million B/D, Hydrotreating Surges,” Oil and Gas Journal, Dec. 23, 1996, pp. 41–48. 9 Krambuhl, C. J., “Gasoline & Diesel Fuel–Where We Are and Where We’re Going,” NPRA Annual Meeting, March 17-19, 1996. 10 Kumar, R., T. S. Thorat, V. Chithra, V. Rathore, V. Peddy, V. C. Rao and N. V. Choudary, “Processing Opportunity Crude Oils—Catalytic Process for High-acid Crudes,” Hydrocarbon World, No. 4, Vol. 2, pp. 64–68, 2009. 11 Rana, M. S., V. Sa´mano, J. Ancheyta and J. A. I. Diaz, “A review of recent advances on process technologies for upgrading of heavy oils and residua,” Fuel, No. 86, pp. 1216–1231, 2007. 12 U.S. Bureau of Standards, Miscellaneous Publication No. 97.

Prashant Parihar is deputy manager (R&D) with Bharat Petroleum Corp., Ltd, India. He has more than five years of research experience in hydroprocessing and optimization of refining processes. He holds an MS degree in chemical engineering from the Institute of Chemical Technology, Mumbai.

Rajeev Kumar is deputy manager (R&D) with Bharat Petroleum Corp.,Ltd., India. His areas of interests are mainly crude oil processing, refining processes, modeling, simulation and optimization. He also has research interest in process development for biodiesel and biolubricants. Mr. Kumar holds an MS degree in chemical engineering from Indian Institute of Technology, Kanpur, India.

Ravi K. Voolapalli is chief manager at Corporate R&D Centre, Bharat Petroleum Corporation Ltd., India. He has 21 years of research experience. His areas of interest are refinery processes, coal-to-liquid technologies, modeling, scale-up and optimization. Dr. Voolapalli holds a BTech degree in chemical engineering from Andhra University, Visakhapatnam, an MTech degree in chemical engineering from Indian Institute of Technology, Kanpur and a PhD in chemical engineering from Imperial College of science technology and Medicine, London.

Sandip Agarwal is deputy manager at the Mumbai Refinery, Bharat Petroleum Corporation Ltd., India. He has 12 years of refining experience. His areas of interests are mainly crude distillation unit, diesel hydrotreating, hydrogen and refinery scheduling and planning. He holds a BS degree in chemical engineering from Laxminarayan Institute of Technology, Nagpur, India. HYDROCARBON PROCESSING MARCH 2012

I 65

Sulfur Experience For the past four decades, Linde Process Plants, Inc. (LPP) has designed and/or constructed sulfur recovery systems. These systems process off gases from gas treating units and sour water strippers. LPP is a world leader in plant modularization and on-site constructed plants. Advantages LPP has the experience, specialized skills, impressive track record, and extensive fabrication facilities to successfully deliver major turnkey process units to global markets. - Extensive technology portfolio - Full turnkey capabilities, including process design, detail engineering, fabrication, construction, and start-up - Competitive pricing - ISO 9001 certified - Single source provider

A member of The Linde Group Linde Process Plants, Inc. 6100 South Yale Avenue, Suite 1200, Tulsa, Oklahoma 74136, USA Phone: +1.918.477.1200, Fax: +1.918.477.1100, www.LPPUSA.com, e-mail: sales@LPPUSA.com Select 85 at www.HydrocarbonProcessing.com/RS

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