Hydrocarbon Processing 03 2016(2).pdf

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MARCH 2016 | HydrocarbonProcessing.com

CORROSION CONTROL Improve corrosion resistance with updated amine treating

MAINTENANCE AND RELIABILITY Detect and prevent boiler leaks in H2 plants using a temperature profiling approach

HEAT TRANSFER Fired heater case study examines high thermal efficiency, low emissions

ENVIRONMENT AND SAFETY Design and install an enclosed ground flare

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MARCH 2016 | Volume 95 Number 3 HydrocarbonProcessing.com

10

38 SPECIAL REPORT: CORROSION CONTROL 39

Improved corrosion prevention with acid-aided regeneration technology D. Lee, J. Klinkenbijl, T. Brok, J. Critchfield and D. Valenzuela

REGIONAL REPORT 45 Central American nations beef up import infrastructure, fuel production amid demand shift M. Rhodes and M. Nogarin

HEAT TRANSFER 53 Calculate thermal efficiency to optimize fired heater operation V. D. Shirpurkar and M. E. Ibrahim

57

Minimize unplanned shutdowns of fired heater operations

DEPARTMENTS 4 10 21 23 107 108 109 110

K. R. Ramakumar

FLUID FLOW AND ROTATING EQUIPMENT 67 Limit the rate of change of fuel gas properties with mixing drum H. Pandya and A. M. Fantolini

NEW IN CATALYSTS—SUPPLEMENT C-73 Taking bio-R&D to commercialization through partnering

M. Carugo and P. Truesdale

101

Design and implement a totally enclosed ground flare M. Choroszy, A. Bourji and P. Prather

PROCESS CONTROL AND INSTRUMENTATION 105 Linear position sensors gain preference in industrial process control applications E. Otto Cover Image: MOL Group’s Danube refinery, located near Budapest, Hungary, has a capacity of 165 Mbpd. It is one of the largest refineries in the Central and Eastern European region.

Industry Metrics Global Project Data Marketplace Advertiser Index Events People

Corrosion prevention is a necessary cost

25

Reliability

27

Automation Strategies

29

Project Management

35

Petrochemicals

37

Engineering Case Histories

D. Sudolsky, J.-P. Burzynski and J.-L. Nocca

ENVIRONMENT AND SAFETY 95 Meet EPA Tier 3 clean fuel regulations through improved blending processes

Business Trends

COLUMNS 9 Editorial Comment

C. Baukal, B. Johnson and R. Newnham

MAINTENANCE AND RELIABILITY 63 Detect boiler leaks upstream of the shift reactor in H2 plants

Industry Perspectives

Avoid pump shaft failures Data analytics solutions require valid data Protect operating margins with outsourced asset management Investigation into West Fertilizer blast shows room for safety improvement Case 89: Cracking of welds due to weld fatigue

www.HydrocarbonProcessing.com

Industry Perspectives Political analysts shed light on US election implications for the downstream In a presidential election year in the US, the hot topic on the minds of many industry leaders in 2016 is the impact those election results might have on the downstream sector. As a result, it comes as little surprise that the majority of the headline speakers at this month’s Annual Meeting of the American Fuel and Petrochemical Manufacturers (AFPM) come from the political realm. Meeting details. The event (FIG. 1), which brings together top

executives and leading personnel from across the entire spectrum of the US downstream industry, kicks off on Monday, March 14 with a general session featuring Mark Halperin and John Heilemann, the managing editors of Bloomberg Politics and co-hosts of the “With All Due Respect” television show on Bloomberg and MSNBC. “With unmatched insider access and keen-eyed perspectives, Halperin and Heilemann will provide an unvarnished take on the headlines and the broader forces shaping American politics,” AFPM said in a statement. The Annual Meeting will also include a special breakfast session on government relations and a speech at the annual luncheon from retired US Gen. Colin Powell. A year ago, those sessions were led by an address from former Hewlett-Packard CEO Carly Fiorina, who dropped hints to industry attendees that she would likely make a run for the US presidency in 2016. A few weeks later, she did exactly that. Full event coverage. As the exclusive show daily provider for

the Annual Meeting, Hydrocarbon Processing will be live at the meeting in San Francisco with full coverage of all conference presentations. The HPInformer blog will include PDFs to the official conference newspapers, as well as pictures and separate news stories regarding all of the meeting’s technical content. Stick with HydrocarbonProcessing.com throughout March for updates from AFPM, as well as many other prominent downstream industry gatherings.

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SUBSCRIPTIONS Subscription price (includes both print and digital versions): Print—One year $239, two years $419, three years $539. Digital format—One year $239. Airmail rate outside North America $175 additional a year. Single copies $35, prepaid. 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.

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For more information about article reprints, call Rhonda Brown with Foster Printing Company at +1 (866) 879-9144 ext. 194 or e-mail [email protected]. 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 © 2016 by Gulf Publishing Company. 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.

EUROMONEY INSTITUTIONAL INVESTOR PLC Directors: John Botts (Chairman), Andrew Rashbass (CEO), Sir Patrick Sergeant, The Viscount Rothermere, Colin Jones, Martin Morgan, David Pritchard, Andrew Ballingal, Tristan Hillgarth Part of Euromoney Institutional Investor PLC. Other energy group titles include: World Oil and Petroleum Economist.

FIG. 1. AFPM’s Annual Meeting typically draws a large downstream crowd, including at last year’s event in San Antonio, Texas.

4 MARCH 2016 | HydrocarbonProcessing.com

President/CEO Vice President, Downstream and Midstream Vice President Vice President, Production Business Finance Manager Publication Agreement Number 40034765

John Royall Bret Ronk Ron Higgins Sheryl Stone Pamela Harvey Printed in USA

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April 5-6, 2016 | Norris Conference Centers - CityCentre

SCM Strategies + Best Practices in a Low-Cost Environment In today’s low-price environment, responding to changing demand with the right capacity, quickly, efficiently and cost-effectively, can make or break profitability. This creates even more imperative for oil and gas companies to build more efficient, cost-effective oil and gas supply chains. At the inaugural O&G Supply Chain Forum (OGSC), you’ll hear from supply chain management experts regarding the latest strategies, trends, and best practices in SCM.

Sessions focus on: • Driving operational excellence thru strategic distribution, planning and scheduling • Cost management • Innovation in technology and research for SCM excellence • Best practices • Developing effective collaborations for an efficient supply chain ecosystem • Efficiencies in shale play operations

Attendees will get a better understanding of: • The framework for building a contemporary and global supply chain: the people, process, product and technology aspects • Current challenges faced by manufacturers and distributors in an environment of lowest total cost • How to use “Big Data” to drive sustainable cost-reduction • New ways of contracting and working internally to deliver substantial cost reductions • 5 key collaborations essential to efficient SCM • The short and long term benefits of category based procurement • The role supply chain solutions play in sand logistics • And much more Lanyard Sponsor:

Exhibitors:

OGSupplyChain.com

Keynote: Craig Freking Vice President Manufacturing and Supply Chain Weir Oil and Gas

Five Key Collaborations Essential to Efficient Supply Chain Management Just as secure supply lines are a key element in military campaigns, effective and efficient supply chain management is critical to success in our global economy. Collaboration with numerous internal and external links in the chain is especially important to manufacturers and distributors operating on the international stage. Freking has 20 years of experience in manufacturing, supply chain management and continuous improvement in a variety of industrial companies including Exterran LLC, Goodyear Tire and Rubber Company, Danaher Corporation and Siemens Stromberg-Carlson. A veteran of the US Marine Corps, he holds a bachelor’s degree in management science from St. Cloud State University and an MBA from Rollins College. He is certified in production and inventory management by APICS, the premier professional association for supply chain management and the leading provider of research, education and certification programs in that field. This presentation will focus on the top five collaboration areas that are critical to successful supply chain management, using examples gleaned from his company’s experience as a global manufacturer of wellhead products for a variety of applications and pressure ratings.

Get valuable insight from these experienced supply chain experts: Capturing cost efficiencies in the market

Implementing category based procurement

Lea Souliotis Global Wells Procurement/ Supply Chain Manager Alaska

Sandeep Singh Manager, Procurement Excellence and Process, Procurement and SCM

BP

Cairn India Limited Soft money approach in bid evaluation

Steve Martin VP – Operations

Saleh A. Hassoubah GM Contracts, Purchasing and Warehouse Dept

Ryder Dedicated East

Saudi Aramco Mobil Refinery Co Ltd (SAMREF) Pioneering new uses of “Big Data” to drive substantial and sustainable cost-reduction

“Breaking new ground” – A supply chain transformation Pravin Tampi VP – Global Sourcing of Supply Chain

Paul Smith VP – Oil and Gas

Newpark Drilling Fluids

Power Advocate

Risk mitigation for efficient capital deployment

Julian Flores Head of Energy Solutions

Douglas Polk VP – Industry Affairs

Panalpina

Vallourec USA Corp Five key trends of oil & gas consolidation

Efficiencies in shale play operations Mark Sen Gupta Senior Consultant

Uday Turaga CEO

ARC Advisory Group

ADI Analytics LLC

>> Download the complete agenda at OGSupplyChain.com For more information about the event / speaking opportunities: contact Tranessa Hunt, Events Coordinator at +1 (713) 520–4470 or [email protected] For sponsor and exhibit opportunities: contact Melissa Smith, Events Director at +1 (713) 520-4475 or [email protected]

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Editorial Comment

ADRIENNE BLUME, EXECUTIVE EDITOR [email protected]

Corrosion prevention is a necessary cost Corrosion is a major maintenance and reliability concern because it has the potential to impact not only plant operations and costs, but also the environment and worker health and safety. It starts as a small problem that can quickly turn into several larger ones. Equipment failures, leaks, plant and unit shutdowns, environmental and product contamination, and worker accidents are a few examples of bigger problems that can occur as a result of corrosion. Given the wide and far-reaching scope of these issues, it comes as no surprise that failure to prevent or mitigate corrosion in critical areas translates into corresponding high costs. Recent incidents emphasize risk. Two high-profile corrosion-related incidents have made headlines in recent months in the US. In late January, a natural gas storage facility near Los Angeles, California was discovered to have released over 150 million tons of methane into the atmosphere since late October 2015. The cement casing of the underground portion of the facility was discovered to be significantly corroded, amid other operational, safety and equipment issues. The incident caused health and relocation problems for thousands of residents. It also required the drilling of a relief well to mitigate the leak. Outside of oil and gas, a public health emergency occurred in Flint, Michigan, when lead-tainted water filtered into the city’s drinking water supply. The Michigan Department of Environmental Quality (MDEQ) decided not to require corrosion control treatment for Flint’s switchover from the Detroit water system to the Flint River in 2014. This caused iron- and leadlined service pipelines to corrode and leach unsafe levels of lead into Flint’s drinking water, as discovered in August 2015. The high costs of corrosion. According to research by inspecting and consulting company G2MT Laboratories, cor-

rosion will cost the US economy more than $1.1 trillion in 2016. This estimate, based on data provided by NACE International, includes direct (operator/owner) costs of corrosion, as well as indirect (non-operator/owner) costs. Design, manufacturing, construction and management costs are included for operators, while non-operator costs may encompass penalties, litigation, environmental cleanup, medical treatment and services suspension. Indirect costs may become direct costs as operators assume responsibility for corrosion-related incidents, when applicable. Among direct costs, money spent on corrosion-prevention measures is generally considered to be money well spent. Such measures include, but are not limited to, the use of specialty coatings, sealants, inhibitors and other protection products; proper materials selection; and regular maintenance, inspection, repair and replacement of corroded equipment and corrosion-prone areas. In the case of the gas storage facility in California, safety and operational risks by the operator combined with undetected corrosion to allow an extended methane release. This release, in turn, translated into mounting public health care, industrial and litigation costs for the company and the state, as well as damaged public perception for the operator. In Flint, the MDEQ’s decision to omit corrosionprevention measures was made in an effort to cut costs. However, this decision ultimately resulted in much higher public health, environmental and repair costs than if corrosion-prevention measures had been installed initially. Both incidents show that prevention and mitigation are the most important costs related to corrosion, as well as the primary tools for addressing it. If applied properly, corrosion-prevention and corrosion-mitigation measures will reduce most other direct and indirect costs associated with corrosion.

INSIDE THIS ISSUE

38 Special Report.

A commonly applied process for the removal of acid gas contaminants from gas is amine treating. At its heart lie the absorption of acid gases into an amine solution and the regeneration of this solvent to be fed back to the absorption column. The interplay between these two steps is important, as the treating performance can be majorly influenced by the amount of acid gas dissolved in the lean solvent returning from the regenerator. Shell Global Solutions discusses corrosion prevention improvements with acidaided regeneration technology.

45 Regional Report.

Due to the growth in the region’s middle class, Central America has seen tremendous petroleum product demand growth over the past decade. This month’s regional report dives into an overview of the region and the future of Central America’s downstream hydrocarbon processing capacity buildout.

57 Heat Transfer.

Four rules to minimize unplanned shutdowns of fired heater operations are shown, as well as examples of potential negative consequences of not following each rule.

105 Instrumentation.

Today’s industrial process control applications increasingly use automated systems to optimize operations and ensure a safer, more productive process. Linear position sensors used in these automated systems provide machine controllers with highly accurate feedback on product parameters, control states and outputs. Eileen Otto discusses how linear position sensors have become the preferred technology for critical and reliable linear displacement measurements in an array of industrial process control applications. Hydrocarbon Processing | MARCH 2016 9

| Business Trends This month’s Business Trends focuses on five major trends affecting the downstream industry. These include: a forecast on global demand for industrial valves to 2020, how the approval of Chinese teapot refinery crude import licenses could lead to a fuel supply glut, the surge in US ammonia and urea plant construction, the start of Nigeria’s DSDP program and the consolidation of Japan's refining industry. Photo: Iowa Fertilizer, a wholly owned subsidiary of Netherlands-based OCI N.V., is building a world-scale fertilizer plant in Wever, Iowa. Once completed, the $1.9-B facility will be able to produce between 1.5 MMtpy and 2 MMtpy of nitrogen fertilizer. Construction on the plant began in 1Q 2013 and is scheduled to begin operations in early 2016. Photo courtesy of Iowa Fertilizer.

LEE NICHOLS, EDITOR/ASSOCIATE PUBLISHER [email protected]

Business Trends World valve demand to reach nearly $100 B by 2020 The Freedonia Group has forecast that world demand for industrial valves will climb to $98.5 B through 2019 (TABLE 1). That represents an annual increase of 4.3%. The company’s latest forecast is detailed in its World Industrial Valves study. The industrial valve market’s greatest growth is in the developing regions of Asia, Africa, Middle East, Central and South America, and Eastern Europe. Out of these regions, China will post the strongest value growth, accounting for 23% of all additional valves sales on a global basis. A number of smaller national markets, such as Indonesia, Malaysia, Thailand and Turkey, will see a healthy increase, as well. Although the US, Western Europe and Japan will see growth in valve demand, they will lag behind the demand growth of developing nations. This is due to below-average increases in process manufacturing output and associated fixed investment expenditures. Market advances will be driven by growth in chemicals and other process manufacturing output, electric power generation and construction activity. Ongoing efforts to expand water infrastructures in developing countries and maintain water and wastewater treatment and distribution systems in developed nations will also contribute to sales increases. The report goes on to say that dollar gains will be boosted by greater use of “smart” valves and actuators, and other better-performing, higher-priced items. The demand for automatic valves (including control and regulator valves, as well as separately sold valve actuators) is projected to grow at a faster pace through 2019 than sales of standard (conventional) hand-operated valves. Automatic valves will continue to take market share away from standard valves because of the advantages they offer, which include improved safety and reduced operating costs. The fastest sales gains of any major product type will be posted by automatic actuators, fueled by valve users’ ongoing efforts to automate standard valve operation by installing automatic actuators as a less costly alternative to replacing units with automatic valves.

Will Chinese teapots worsen China’s fuel glut? In 2015, China loosened restrictions on independent refiners’ ability to secure crude oil from the international market. These refiners, known as teapots, have a capacity of 20 Mbpd to 100 Mbpd. Although these refineries tend to be less complex than their nationally-owned counterparts, teapot refineries account for one-third of China’s total domestic refining capacity. China’s total refining capacity topped 14 MMbpd in 2015. Also within that year, expansions in teapot refining operations increased the independent refining sector’s total refin-

ing capacity to nearly 4.5 MMbpd, or nearly one-third of total domestic refining capacity. Allowing teapot refineries to access international crude is another step being taken by the Chinese government to deregulate its oil market. China’s administration, headed by President Xi Jinping, wants market forces to play a more decisive role in the industry. Allowing additional crude oil imports from small, independent refiners will increase competition, as well as allow additional entrants into the industry. For 2016, the Chinese government has more than doubled non-state crude import quotas to nearly 88 MMtpy, or approximately 1.75 MMbpd. The new import licenses have restrictions, though, that include: • The refiner must have one crude distillation unit with processing capacity of at least 2 MMtpy (40 Mbpd) • The refiner must have an available credit line of $1 B certified by commercial banks • The company must have at least five people with over five years of international trading experience • The company must have at least 300 Mt of oil storage capacity • The refiner must meet certain environmental conditions, which include the decommissioning of older, morepolluting units. The first import license was awarded in early August 2015 to independent Chinese refiner Baota Petrochemical Group. By mid-August, two additional Chinese independent refiners received import licenses. These were Shandong Dongming Petrochemical Group and Panjin Beifang Asphalt Fuel Co. By the end of September, Linjin Petrochemical Plant Co. and Shandong Kenli Petrochemical Group were also awarded crude import licenses. At the time of this publication, additional crude import licenses have been awarded to: • Shandong Huifeng Petrochemical Group • Chambroad Petrochemicals Co. TABLE 1. World industrial valve demand forecast through 2019 Total industrial valve demand, US $ MM

Annual growth, %

2009

2014

2019

14,300

19,300

22,650

6.2

3.3

Western Europe 13,400

15,800

18,100

3.3

2.8

Asia-Pacific

19,640

26,900

35,250

6.5

5.6

Central and South America

3,340

4,650

5,870

6.8

4.8

Eastern Europe

4,710

6,450

7,800

6.5

3.9

Africa/ Middle East

4,810

6,800

8,830

7.2

5.4

5.8

4.3

North America

Total

2009–2014 2014–2019

60,200 79,900 98,500

Source: The Freedonia Group’s World Industrial Valves study

Hydrocarbon Processing | MARCH 2016 11

Business Trends • Sinochem Hongrun Petrochemical Corp. • Tianhong Chemical • Shandong Shouguang Luqing Petrochemical Co. In 2015, 10 teapot refineries were awarded crude import licenses totaling over 43 MMtpy. An additional 10 applications, representing over 38 MMtpy, are awaiting government approval. These include: • Dongying Qirun Chemcial Co. • Hebei Xinhai Chemical • Henan Fengli Petrochemical Co. • Shaanxi Yanchang Petroleum Group • Shandong Haiyou Petrochemical Group • Shandong Hengyuan Petrochemical Co. • Shandong Jincheng Petrochemical • Sinochem Hongrun Petrochemical Co. • Shandong Qingyuan Group • Wudi Xinyue Ran Hua Co. If approved, total crude imports by Chinese teapots could surpass 82 MMtpy, or nearly 1.7 MMbpd, in 2016. The majority of this crude will flow to the Shandong province. Located in northeast China, the province contains approximately 80% of China’s teapot refineries. The province has already received license approvals for nearly 31 MMtpy of crude imports, with an additional 26.8 MMtpy awaiting approval. An additional 25 MMtpy of crude import licenses awaiting approval are located in the Liaoning, Ningxia, Henan, Hebei and Shaanxi provinces (FIG. 1). Should all crude import licenses be approved, China could overtake the US to become the world’s largest crude oil importer. The big question is how will the increased market share of crude oil processing for China’s teapot refineries affect domestic output and the region as a whole? Now that Chinese independent refiners can utilize crude oil in lieu of low-quality fuel oil, non-state refineries are expected to boost run rates. This, in turn, will create more refined products, which could add to the fuel supply glut already being witnessed domestically and in many countries in the Asia-Pacific region. Since the majority of Chinese independent refiners lack infrastructure to export their

Ningxia Approved: 6.16 MMtpy Shaanxi Waiting: 3.6 MMtpy

Hebei Waiting: 5.4 MMtpy

Henan Waiting: 2.93 MMtpy

Liaoning Approved: 7 MMtpy Shandong Approved: 30.73 MMtpy Waiting: 26.8 MMtpy

FIG. 1. Breakdown on the approval status of Chinese teapot refiners’ crude import licenses by province.

12 MARCH 2016 | HydrocarbonProcessing.com

products to the global market, their refined products will be sold primarily to the domestic market. This will ultimately eat into market share held by state-owned entities China National Petroleum Corp. (CNPC), the parent company of PetroChina; China Petroleum and Chemical Corp. (Sinopec); and China National Offshore Oil Corp. (CNOOC). Chinese teapots’ growing crude processing market share may force state-owned refiners to either find additional export markets or cut run rates. Large, state-owned refiners are already adding to a diesel supply glut in the region, however. According to China’s National Bureau of Statistics, domestic refining output increased by nearly 4% in 2015, reaching nearly 10.5 MMbpd. The increase in crude oil processing was in response to a surge in domestic demand for gasoline. Chinese refineries were built mainly to satisfy domestic diesel demand. However, overcapacity and slowing industrial buildout have created an oversupply of diesel, which led the country to become a net diesel exporter in 2014. China’s shift to a more consumer- and service-based economy has lessened the demand for diesel from the construction and heavy-duty trucks sector. As China’s middle class expands and car sales rise, additional gasoline demand has caused domestic refiners to produce more gasoline supplies. In doing so, these refineries ultimately produce additional diesel supplies. Now that the country’s industrial growth is slowing, excess Chinese diesel supplies could end up flooding the Asian market. With Chinese teapots likely to increase refinery run rates, state-owned refiners will have to combat independents eating into their domestic market share. Should state-owned refiners be unable to find export markets for excess diesel supplies, and both independent and state-owned refiners neglect to cut refinery run rates, then fuel stocks will inevitably begin to rise, and the nation’s fuel glut will continue to worsen.

US 2016, the year of ammonia-urea plant capacity growth It is not breaking news that the US shale gas boom has sparked a surge in the construction of new processing capacity. Some of the largest capital-intensive investments are being made in the US petrochemical and gas processing/LNG sectors. Total announced capital investments in the US petrochemical sector have eclipsed $135 B. New capacity includes a sharp increase in the construction of ethane cracking and derivatives capacity, methanol plant construction and propane dehydrogenation units. The availability of cheap shale gas is also expanding the US fertilizer industry. Over $16 B in new ammonia-urea plant projects have been announced. The majority of these projects are located in the Midwest, near agricultural demand centers. FIG. 2 provides the location for active ammonia and urea projects in the US. Total capital expenditures have decreased from over $20 B. In 2015, CHS abandoned its $3-B Spiritwood project in North Dakota, and EuroChem scrapped its $1.5-B Louisiana ammonia plant. CHS announced that its Spiritwood project was abandoned due to increasing construction costs, water supply challenges, overall risk profile and time required for the project to be built. In the end, CHS decided that the return on investment could not justify the project’s cost. EuroChem shelved its $1.5-B Louisiana project due to difficulty in securing financing. Also, Western sanctions against Russia closed off many financ-

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Business Trends ing options for the company. EuroChem decided to abandon the 800-Mtpy project in the second half of 2015. Regardless, the US will see a hefty buildup of ammonia-urea capacity through the end of the decade. Over 5 MMtpy of additional capacity is scheduled to go online in 2016. This includes capital-intensive projects such as CF Industries’ Donaldsonville, Louisiana plant (began full-ramp up operations in 4Q 2015) and its Port Neal, Iowa plant, and Dyno Nobel’s Waggaman project. Total new ammonia-urea capacity could top 10 MMtpy by 2020 if all projects are completed. Northern Plains Nitrogen Dakota Gasification Magnida

JR Simplot

Grannus

Total ammonia-urea capacity startup 2016 = 5.5 MMtpy 2017 = 4.3 MMtpy 2018 = 2.5 MMtpy

OCI CF Industries

Cronus Chemicals Ohio Valley Midwest Fertilizer Resources Koch Austin Powder Koch LSB Industries Southern Co. (ammonia Gulf Coast byproduct) Ammonia (no site yet) Pallas AM Agigren BASF-Yara Nitrogen CF Industries Dyno Nobel

FIG. 2. Active ammonia-urea plant projects in the US. Source: Hydrocarbon Processing’s Construction Boxscore Database.

Nigeria begins DSDP program to aid transparency Beginning this month, Nigeria will replace its oil swaps scheme, also known as offshore processing arrangements, with a direct-sale-direct-purchase (DSDP) program. The new DSDP program will allow state-owned Nigerian National Petroleum Co. (NNPC) to sell crude oil and buy refined products directly from international refineries. The new program is intended to end the country’s oil swaps program, as well as save the country millions of dollars in costs. The oil swaps program was started in 2010 when the cashstrapped NNPC could no longer pay for imported gasoline and diesel with cash. The oil swaps program was basically an oil bartering scheme that utilized a middleman to pay for refined products from foreign partners. The Nigerian government provided crude oil to traders in exchange for refined imports, such as gasoline and diesel. It is believed that these types of contracts lacked transparency and may have cost the Nigerian government billions of dollars over the past five years. The Nigerian government has decided to combat unreliable oil swaps with its DSDP program. The DSDP program’s goal is to provide a more transparent system of crude oil and refined product purchases, as well as save the Nigerian government billions of dollars by eliminating the cost of using a middleman. If the DSDP initiative is successful in saving the Nigerian government billions of dollars, then the savings could be used

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Business Trends to finally update the country’s downstream processing sector. The majority of downstream activities in the country are located in the Niger Delta. The area contains three of the country’s four operating refineries. Nigeria’s four refineries (Port Harcourt I and II, Warri, and Kaduna) have a total operating capacity of 450 Mbpd. This refining capacity, which is frequently underutilized, is not enough to meet domestic demand. The country’s crude oil production is nearly eight times higher than domestic consumption (FIG. 3), but poor refinery utilization forces the country to rely on refined fuel imports to satisfy demand. Numerous refinery projects have been announced in recent years, but few have come to fruition. Projects. Although the country has witnessed its fair share

of project holds and cancellations, multiple downstream projects are likely to move ahead. The most notable project is the construction of Africa’s largest privately owned refinery. The Dangote Industries Ltd. (DIL) integrated complex will be constructed in Lekki, Lagos State, Nigeria, and will include a petrochemical complex and fertilizer facility (FIG. 4). The project will be the first of its kind in Nigeria. The $9-B, 650-Mbpd refinery will include a petrochemical plant that will produce 750 Mtpy of polypropylene, and a fertilizer plant that will produce 2.8 MMtpy of urea and ammonia for the nation’s agriculture sector. The refinery will produce Petroleum and other liquids production and consumption in Nigeria, MMbpd

3.0

Total oil production

2.5

Crude oil production

2.0 1.5 1.0

Crude output falls by more than 25% from 2005 to 2009 as infrastructure attacks and oil theft escalate.

Net exports

0.5

0.0

Total oil consumption 2005

2006

2007

2008

2009

2010

Attacks on oil facilities declined following the implementation of the amnesty program (2009–2010). However, oil production has been stagnant or declining over the past few years because of supply disruptions and natural production declines.

2011

2012

2013

2014

FIG. 3. Petroleum and other liquids production and consumption in Nigeria. Source: US EIA.

gasoline, diesel, aviation fuel and slurry to be used as a raw material for carbon black. The refinery’s primary goal is to supply the local market and reduce refined fuel imports. The project was also a nominee for Hydrocarbon Processing’s 2015 Top Project awards for refining. Completion is scheduled for 2018. The announcement of the DIL refining and petrochemical complex has motivated additional companies to pursue largescale projects. Brass Fertilizer is planning to construct a worldscale methanol, ammonia, urea granulation and gas processing plant on Brass Island. The $3.5-B Brass Fertilizer project will produce 5 Mtpd of methanol, 2.2 Mtpd of ammonia and 7.7 Mtpd of urea for domestic and export markets. Operations are scheduled to begin in 2018. Indorama is expected to ramp up its $1.8-B Eleme fertilizer plant to full commercial operations in 1Q 2016. The 1.4-MMtpy single-train urea plant is the largest in Nigeria. The project was created in response to the Nigerian government’s plan to privatize the country’s fertilizer industry. The gas-to-ureabased plant is part of Indorama’s aim to create the continent’s largest petrochemical hub. The complex will ultimately reduce urea imports, as well as provide affordable nitrogen-based fertilizers to the growing agriculture industry. Quantum Petrochemical recently built a grassroots petrochemical and methanol complex. The $1.5-B facility is located in southern Nigeria on the Gulf of Guinea. The facility will produce ethylene derivatives such as polyethylene, polypropylene and methanol. Gulf of Guinea Methanol Ltd., a subsidiary of Nigeria’s Gulf of Guinea Oil Exploration Ltd., is still looking into developing a $1.1-B plant to convert natural gas into methanol. The methanol produced would be used as a feedstock for the petrochemical industry. Lastly, Nigeria LNG is still committed to building a seventh train at its Bonny Island LNG terminal. Construction of Train 7 would cost approximately $2.5 B. The project has been in limbo for numerous years, and no date for a final investment decision has been announced. If built, Train 7 would raise total capacity at Bonny Island from 22 MMtpy to 30 MMtpy.

FIG. 4. Construction continues on the DIL refinery and petrochemical integrated complex in Nigeria. Photo courtesy of the Dangote Group.

16 MARCH 2016 | HydrocarbonProcessing.com

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Business Trends

The island nation is embarking on a massive reduction in refining capacity. Oil products consumption in Japan has decreased for over a decade due to a shrinking population that is transitioning to more efficient hybrid vehicles for transportation. The substantial decrease in demand for transportation fuels has resulted in excess domestic refining capacity. In turn, refinery utilization has been reduced or eliminated. According to BP’s Statistical Review of World Energy 2015, recent Japanese refining capacity peaked at 4.65 MMbpd in 2008, and has been in decline ever since (FIG. 5). By 2015, Japan’s refining capacity had dropped over 900 Mbpd to nearly 3.75 MMbpd. On top of declining demand, the Japanese government is seeking to promote operational efficiency through its Refining Ordinance. The plan was introduced in 2010 by the Japanese Ministry of Economy, Trade and Industry (METI), and called for a new mandatory cracking-to-crude distillation capacity ratio of 13%. The new rules became effective in March 2014. The country’s second phase of its Refining Ordinance is likely to shed up to 400 Mbpd of additional domestic refining capacity by 2017. To adhere to the mandatory requirements, Japanese refiners are expected to decrease utilization rates, consolidate operations or shut down facilities. By the end of 2015, four of the five largest refining companies had already announced mergers. Idemitsu Kosan will take over Showa

Shell Sekiyu, and JX Holdings, Japan’s largest refiner, will merge with TonenGeneral Sekiyu. These mergers are expected to take place by the end of 2Q 2017, and will result in two firms dominating the Japanese domestic fuels market. With declining domestic fuels demand, the METI’s refining ordinances and domestic refiners merging operations, Reuters’ analysis forecasts that Japanese refining capacity could decrease to 3.2 MMbpd by 2020, and down to 2.3 MMbpd by 2030. If this scenario plays out, Japan would be dependent on refined fuels imports to meet demand, especially for gasoline. 6 Refining capacity Consumption

5

Forecast

4 MMbpd

The future of Japan’s refining industry

3 2 1 0

2005 2006 2007 2008 2009

2010

2011

2012

2013

2014

2020 2030

FIG. 5. Oil consumption vs. refining capacity in Japan with refining capacity forecast to 2030. Source: BP Statistical Review of World Energy 2015, Reuters’ analysis.

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UOP delivers complete modular process units for the petroleum refining, petrochemicals and gas industries. Modular delivery minimizes overall project schedule, cost and risk by maximizing prefabrication of complex process units under quality controlled conditions and under the watchful eye of the process licensor. UOP has delivered more than 1,400 fully engineered and fabricated UOP process units to the global oil and gas industries.

For more information about UOP modular solutions, visit www.uop.com © 2016 Honeywell International. All rights reserved

MIKE RHODES, MANAGING EDITOR [email protected]

Industry Metrics

15 10

7

Cracking spread, US$/bbl

Feb.-16

Dec.-15

Nov.-15

Oct.-15

Sept.-15

Aug.-15

July-15

June-15

May-15

April- 15

Mar.-15

Feb.-15

Jan.-16

Jan.-16

Dec.-15

Nov.-15

Oct.-15

Sept.-15

Aug.-15

June-15

July-15

Jan.-16

Dec.-15

Nov.-15

Oct.-15

Sept.-15

Aug.-15

July-15

Jan.-16

Dec.-15

Nov.-15

Oct.-15

Sept.-15

Aug.-15

July-15

June-15

May-15

April-15

Jan.-16

Dec.-15

Nov.-15

Oct.-15

Sept.-15

Aug.-15

July-15

June-15

Jan.-16

Dec.-15

Nov.-15

Oct.-15

Sept.-15

Gasoil Fuel oil

Aug.-15

Dubai Urals

Prem. gasoline Jet/kero

July-15

0

0

-10 -20

June-15

2

10

Jan.-15

4

20

May-15

Cracking spread, US$/bbl

30

6

Jan.-15

May-15

Singapore cracking spread vs. Oman, 2015–2016*

Brent dated vs. sour grades (Urals and Dubai) spread, 2015–2016* Light sweet/medium sour crude spread, US$/bbl

Gasoil Fuel oil

-10 -20

Source: EIA Short-Term Energy Outlook, February 2016

-2 -4

Prem. gasoline Jet/kero

April-15

2017-Q1

0

April-15

2016-Q1

10

Mar.-15

2015-Q1

30 20

Mar.-15

2014-Q1

Stock change and balance, MMbpd

Supply and demand, MMbpd

6 5 4 3 2 1 0 -1 -2 -3

Forecast

40

Jan.-15

2013-Q1

Feb.-15

Rotterdam cracking spread vs. Brent, 2015–2016*

World liquid fuel supply and demand, MMbpd Stock change and balance World supply World demand

Mar.-15

J F M A M J J A S O N D J F M A M J J A S O N D J 2014 2015 2016

2012-Q1

Prem. gasoline Jet/kero Diesel Fuel oil

Cracking spread, US$/bbl

Source: DOE

Jan.-15

Oil prices, $/bbl

60 50 40 30 20 10 0 -10 -20

W. Texas Inter. Brent Blend Dubai Fateh

100 98 96 94 92 90 88 86 84 82 2011-Q1

April-15

US Gulf cracking spread vs. WTI, 2015–2016*

Selected world oil prices, $/bbl 120 110 100 90 80 70 60 50 40 30 20

June-15

Production equals US marketed production, wet gas. Source: EIA.

Japan Singapore

May-15

60

US EU 16

April-15

J F M A M J J A S O N D J F M A M J J A S O N D J 2014 2015 2016

70

Feb.-15

0

80

Feb.-15

20

2 1 0

Mar.-15

3 Monthly price (Henry Hub) 12-month price avg. Production

90

Feb.-15

4

Utilization rates, %

60

100

Jan.-15

5

Gas prices, $/Mcf

Production, Bcfd

Global refining utilization rates, 2015–2016*

6

80

Mar.-15

US gas production (Bcfd) and prices ($/Mcf)

Feb.-15

Jan.-15

0

100

40

WTI, US Gulf Brent, Rotterdam Oman, Singapore

4

May-15

An expanded version of Industry Metrics can be found online at HydrocarbonProcessing.com.

Global refining margins, 2015–2016* 20

Margins, US$/bbl

Refinery margins weakened in the US due to lower winter demand. In Europe and Asia, stronger regional demand amid a tightening environment and export opportunities allowed for a recovery in crack spreads at the bottom of the barrel, which, along with continued strength in the gasoline market, allowed refinery margins to rise. Refinery utilization rates rose, mainly in Asia, following the end of the heavy maintenance season.

* Material published permission of the OPEC Secretariat; copyright 2016; all rights reserved; OPEC Monthly Oil Market Report, February 2016. Hydrocarbon Processing | MARCH 2016 21

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LEE NICHOLS, EDITOR/ASSOCIATE PUBLISHER [email protected]

Global Project Data Presently, Hydrocarbon Processing’s Construction Boxscore Database is tracking over 2,100 projects around the world. The map below shows a breakdown of total active downstream projects by region and status. At the time of this publication, approximately 60% of active projects are in the preconstruction

stage. The Asia-Pacific region continues to lead in total active projects, followed closely by the Middle East. Approximately 40% of active projects in both regions are under construction. New project announcements increased in February, primarily from India and the US.

21

20 9 9

80 65

65

4

Canada

96

72

8 29

19 150

Europe

37

12

US

33

30

87

13 14 16 47

41

Planning Study Feed Engineering Under construction

14

Africa

20

79

60

59

16

185

165

Middle East 46

72 22

Asia-Pacific

Latin America

Total active projects by region and status, March 2016 30 24 21

26

25

27 22 17

18

27

26 20

18

6% Other 36% Refining

13

33% Petrochemicals Jan.- Feb.- Mar.- April- May- June- July- Aug- Sept.- Oct.- Nov.- Dec.- Jan.- Feb.15 15 15 15 15 15 15 15 15 15 15 15 16 16

Boxscore new project announcements, January 2015–present

25% Gas processing/LNG Market share breakdown of downstream HPI projects by sector

Detailed and up-to-date information for active construction projects in the refining, gas processing and petrochemical industries across the globe | ConstructionBoxscore.com Hydrocarbon Processing | MARCH 2016 23

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Reliability

HEINZ P. BLOCH, RELIABILITY/EQUIPMENT EDITOR [email protected]

Avoid pump shaft failures We welcome and thrive on questions from reliability engineers and try to answer some as best we can. Inquiries do not only inform us of the state of knowledge and the issues confronting industry; our answers often require research, and our selections are appreciated by other readers. We were asked to provide some direction relating to a reader’s research on potential causes of catastrophic pump shaft and impeller failure. “In the course of my work as a pump and vibration specialist,” the reader said, “I have encountered a number of such failures. One root cause of these failures that has been suggested is a failed check valve, which allows the pump to freewheel in reverse due to backflow. However, I have found no documentation that supports or explains this failure mode. In conversations with various pump and field service technicians, it has been suggested that, when a pump is started while spinning in reverse, the starting torque exceeds the shaft strength and catastrophic shaft failure occurs. However, this explanation is counter to my understanding of electric motor starting torque.” Pumps and pumps. Our reply highlighted first what our reader, of course, knew: There are cheap pumps and there are welldesigned, more expensive pumps. In the days when there was still an abundance of common sense, some wise man wrote that we always get what we pay for. User experts Ed Nelson and John Dufour1 noted that nearly all impeller fastening arrangements for single-stage pumps are threaded in a direction that is counter to the as-designed pump rotation. This reduces the chances of loosening in case the pump is driven in the reverse direction. But, the probability of reverse rotation is close to zero if the motor’s direction of rotation is checked before coupling the driver to the driven shaft. If the motor and pump are installed in the as-shipped and coupled condition, the chances of an impeller coming off are, of course, 50%.2 Nelson and Dufour remind us that some impellers are screwed onto (or into) the wet end of the pump shaft (FIG. 1). These impeller types are particularly vulnerable to backward rotation in the case of product backflow. Backflow can be prevented by discharge check valves, but not all pumping loops have working check valves, and user companies rarely include these valves on their periodic turnaround inspection and repair schedules. That fact is kept in mind by users who, for that and other reasons, favor pump designs that safely secure impellers to the shaft. Know the impeller fastening method. No pump manufacturer has a universal fastener suitable for all pump sizes and service environments. In fact, the fastening method is not usually shown on the manufacturer’s standard drawings. Also, relatively few user-purchasers include process pumps in a thorough

upfront machinery quality assessment (MQA).3 The fasteners, in general, are standard items that the pump manufacturer purchases in bulk from a competitive seller. Like pump owners, pump manufacturers want to purchase parts from the lowestcost supplier of buy-out or third-party items. However, these fasteners can be a source of problems. Hardness and metallurgy must be observed, which brings us back to an MQA. Usually, an impeller spins off only if it is not properly secured, but even a keyed impeller fit jeopardizes reliability if the key is loosely fitted. Whatever their size and speed, pump impellers secured by castellated nuts or washers must retain the impellers in a manner that does not allow them to come off while operating in any direction. That is why we should examine drawings before we purchase; we should know how parts or machines work before we buy parts; and we must ask a lot of relevant questions before buying plant assets. Some superior pump designs use keys to secure impellers. A good key fit is a snug fit, meaning that hand-fitting is needed. Whereas, a vulnerable keyway has sharp corners, a keyway with low-stress concentration has bottom fillet radii, which then increases shaft safety factors. A well-designed shaft end also has a generous fillet radius at the shaft shoulders. In some applications, the fillet contour is purposely made so large that special care must be taken so that it does not interfere with the mating radius at the bearing inner ring. Verification takes time. The indifferent do not take the extra time, and the uninformed do not even know what to do with the extra time management allocated for higher-quality work. Pardon us for quoting Mark Twain by saying, “The man who has a book and does not read it is no different from the man who cannot read.” For the record. The starting torque of many motors is seven times the full running torque. It is agreed that discharge check

A

FIG. 1. A screw-on style impeller, which is particularly vulnerable to backward rotation in the case of product backflow.2 Hydrocarbon Processing | MARCH 2016 25

Reliability valves rarely leak to the point of allowing substantial reverse flow. “Lean and mean” plants don’t always install these check valves, and, if they do, those valves are not usually included in preventive-maintenance routines. In any event, a pump reliability and/or failure analysis review should include the piping and all related systems. In the reader’s repeat-failure example, it is also possible that several seemingly small deviations can combine. It is easy to get away with one or even two deviations, but it is rare to succeed with four or five. Next to an electric motor, a pump is the simplest machine used by man. It typically has 40 parts and yet fails relatively often. An aircraft jet engine has more than 8,000 parts and rarely fails. Why? Jet engine manufacturers strive for perfection; they disallow every known deviation. Their quest to find root causes of failure and their refusal to tolerate known deviances require trained personnel, strict adherence to checklists and procedures, and the time to do things correctly. In this instance, we were not given enough information to accurately determine why the reader’s pump shafts failed. We can only vouch for the greatly increased probability of combining a few seemingly minor deviations from best practice so as to cause trouble. As deviations combine, safety factors will vanish, impellers will come off and shafts will break. There will never be a good substitute for following procedures and for understanding what happened. Replacing parts and restarting the rebuilt machine without addressing the true root cause is the perfect setup for repeat failures. Finding the root causes of fail-

ure and implementing sound remedial steps are the common sense courses of action. From our response, the reader correctly inferred that verbal hints at failed check valves causing catastrophic shaft failures are not supported by much factual evidence. He found nothing in the physics of how electric motors develop torque that could justify the over-torque reports. He suspected that such an explanation for a pump shaft failure was a convenient, but unverifiable, explanation, when identifying the real cause is somewhat inconvenient. We essentially agreed, but gave him additional food for thought. LITERATURE CITED Dufour, J. W. and W. E. Nelson, Centrifugal pump sourcebook, McGraw-Hill, New York, New York, 1992. 2 Bloch, H. P., Pump wisdom: Problem solving for operators and specialists, John Wiley & Sons, Hoboken, New Jersey, 2011. 3 Bloch, H. P. and A. R. Budris, Pump user’s handbook: Life extension, 4th edition, The Fairmont Press, Lilburn, Georgia, 2013. 1

HEINZ P. BLOCH resides in Westminster, Colorado. His professional career commenced in 1962 and included long-term assignments as Exxon Chemical’s regional machinery specialist for the US. He has authored over 650 publications, among them 19 comprehensive books on practical machinery management, failure analysis, failure avoidance, compressors, steam turbines, pumps, oil-mist lubrication and practical lubrication for industry. Mr. Bloch holds BS and MS degrees in mechanical engineering. He is an ASME life fellow and maintains registration as a professional engineer in New Jersey and Texas.

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26 MARCH 2016 | HydrocarbonProcessing.com

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Automation Strategies

RICK RYS, SENIOR CONSULTANT ARC Advisory Group

Data analytics solutions require valid data Process plants collect enormous amounts of data, both from the sensors connected to the control system and from other sources, including edge devices in the plant and remote devices in the cloud. However, to gain value from all of this data, it must be properly collected, validated, analyzed and visualized. Data quality is particularly important. Data quality status critical. When the first digital con-

trol systems were developed, it became apparent that digital controllers required data with a known quality status. This is because controllers acting on bad data could lead to costly and often perilous process upsets. Distributed control system (DCS) developers embedded the quality information side by side with the data as a quality status to enable the controller to know when a measured value was invalid and act accordingly. However, with many control systems, it is difficult to view the data status, making it a common error to forget about the data status when collecting data in a real-time historian or when using that data in a calculation. Without knowing the difference between valid and invalid historical data, decisions made or actions taken on that data could be flawed.

Time factor also important. The time factor is also important, specifically the sampling frequency and the timestamping of data. The input/output (I/O) system involved with industrial manufacturing gathers sensor data with sample rates typically in the range of 1 millisecond (msec) to 25 msec. Real-time control systems can gather data asynchronously from the I/O subsystem, often with some averaging and signal conditioning involved. The controller typically samples all inputs, executes all control algorithms, and writes all outputs in a fixed control cycle. Modern control systems typically operate at between 10 msec and 1,000 msec. Some data acquisition and control systems timestamp every unique data value. In the manufacturing industry, the sample rate for data in a historian is roughly an order of magnitude slower than the controller cycle. Scan times for data collection are commonly between 1 sec and 60 sec. In the typical configuration, the data historian collects the data at a constant time interval and internally timestamps the data when collected. For many data analytics applications, the I/O and control system delays are so small that they can be ignored. A classic problem is “data aliasing.” Slowly taking regular samples of an oscillating waveform (like a sine wave) results in recording a harmonic. This obscures the actual oscillation. As a rule, it is necessary to sample at least 10 times faster than the time constant of interest in the data analytics application. Research data loggers may need to sample data much faster than typical industrial historians.

In instances where data might be collected from multiple sources with transmission delays, the sensor data can be timestamped at the moment the measurement is made to ensure accurate time synchronization. In some cases, I/O systems compute a running average of several samples, and the control systems collect this averaged value. A common practice is to filter noisy analog data with a first-order lag or similar filter in the control system. In most cases, these digital data manipulations are not harmful, but care should be taken to understand how the data gets from the sensor to the historian to make sure it remains useful. Different approaches used. Distributed control system and programmable logic controller suppliers recognize the need to handle quality status and timestamping, but each supplier solves this problem in its own way, within its own architecture. Combining data from the Industrial Internet of Things (IIoT) with traditional industrial historian data can be problematic, as the pedigree of the data can vary widely and will depend on the communication architecture. Data quality status needs to be built into the sensor, the signal conditioning and the communications technology that moves the data from the sensor into applications like a data historian. A useful approach is to define each data record as either “valid” or “invalid” by inspecting all of the data validity information. If the status is “bad,” “out of service,” “error,” “high out of range,” “low out of range,” or “I/O BLOCK IN MANUAL,” or any other fatal error, then the data should be marked as “invalid.” Anyone looking at historical data is likely to ask, “What was really going on in my process?” during a time period of interest. This requires uniform collection of valid sensor data, along with contextual information that allows sensor and operational data to be sorted. Processes have modes of operation like starting, normal operation, production grade, grade transition, calibrating, stopping and shutdown. It is helpful to compute the context in the control system and to record it as part of the data. The context is like a “treasure map” that can show data analysts how to gain maximum value from their data. RICK RYS, a senior consultant at ARC, performs research and consults with clients on technology areas, such as process automation, energy management, advanced process control (APC), simulation and optimization. Prior to joining ARC, Mr. Rys worked as an independent engineering consultant at R2Controls since 1996. He also worked for Foxboro (Invensys and now Schneider Electric) for 20 years in the process industries. Mr. Rys holds a BS degree in chemical engineering from the University of Massachusetts, and he is a registered professional chemical engineer in Massachusetts. Hydrocarbon Processing | MARCH 2016 27

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Project Management

CHUCK NEWMISTER, ACCOUNT MANAGER SKF, Austin, Texas

Protect operating margins with outsourced asset management The oil and gas sector is under pressure to maintain operating margins. Although oil prices have stabilized in recent months, they have, nonetheless, fallen dramatically in the last year. This has been driven mainly by oversupply, which is due to a range of factors that include new reservoir discoveries and drilling techniques, the shale fracking boom, weakness in major economies, the growing availability of alternative energy sources and the introduction of ever-more-efficient combustion engines in cars, trucks, and industrial and power generation plants. Although the market for natural gas is slightly healthier, there is also growing concern as short-term supply is likely to exceed demand, with industry analysts predicting only a modest annual growth of 2.4% to 2018.1 Continuing volatility across the supply chain seems likely to become the norm in the immediate future. Consequently, companies must prepare for uncertainty and adapt operations to become increasingly agile, while developing improved methods to derive maximum value from existing and future investments in extraction, refining, and distribution systems and equipment. Asset management and cost optimization. From the

early days of volume production and commercialization, the oil and gas sector has led the world in the use of predictive and preventive maintenance, developing models that have subsequently been followed by many other sectors of industry. These techniques have been used as key tools to improve plant and process safety, efficiency and optimization, and have underpinned the growth of a new discipline of asset management. Asset management is based on a strategic assessment that identifies plant improvement opportunities based on criticality, and then defines and applies the most appropriate solutions, which are designed specifically for each business. The objective is to incorporate business goals, application challenges and organizational culture into a road map that improves the reliability, performance and functionality of all operational extraction, process and distribution assets. A methodology of asset efficiency optimization (AEO) is then applied to ensure that every asset is utilized as efficiently as possible, and to maximize output without increasing capital expenditure, all while reducing overall maintenance and operational costs. A successful asset management program depends on a clearly defined strategy that is driven by business goals, beginning with an understanding of the current industry environment and a vision of where the business needs to be to achieve

optimum short- and long-term performance. This can be a difficult process to manage due to the complexity of the production and management systems found in many large, multi-site, multinational oil and gas businesses. Charting a course of action. Generally, the first step is to carry out a client needs analysis (CNA) that is based on a straightforward 40-question survey to provide a snapshot of the operation of each production facility. This provides a measurable overview of the way in which its reliability processes are functioning, as well as its position on the maintenance maturity continuum, benchmarked against industry averages and best practices. Once complete, the CNA provides the key data needed to draw up a detailed AEO plan to improve plant reliability and asset utilization. This work management procedure addresses four key areas: maintenance strategy, work identification, work control and work execution, all of which provide an integrated methodology that reflects the unique processes, culture and technology at each facility or operation (FIG. 1). A structured approach to asset management will be familiar to many companies in the oil and gas sector, especially larger organizations. However, over time, the carefully planned long-term asset management programs become disrupted due to financial restrictions, plant updates, company acquisitions, changes in regulations or the launch of new products. As a result, the viewpoint of managers and engineers becomes focused on short-term challenges and internal issues, pushing longer-term plans down the order of priority. It is also worth noting that, although many companies in the sector have extremely effective asset management and plant reliability processes, these are not always systemized. Information and knowledge that have been gained over many years by plant engineers and operators largely remain locked in their heads, with only a small proportion being recorded and catalogued in a way that can easily be assimilated by new employees or third-party contractors. Consequently, when employees leave, they often take years of ingrained experience and knowledge with them, leaving their replacements struggling to manage the demands of often highly complex process systems. A CNA analysis is a simple method of beginning the process, and can lead to the next phase, a reliability-centered maintenance (RCM) project. This makes it possible to begin capturing much of this valuable information in a way that is meaningful to plant engineers and senior managers alike, and in a format that reflects other business processes and can be used on an ongoing basis. Hydrocarbon Processing | MARCH 2016 29

Project Management Long-term vision. The scale and complexity of most oil and gas operations demand that, to be truly successful, an asset management strategy requires a clear vision and a long-term tactical implementation plan. Anything less will almost certainly lead to an increase in operating costs, with the risk of burgeoning levels of equipment downtime, system reliability and loss of productivity. The growing market volatility and pressure on margins, combined with factors like staff and skills shortages, and the particular demands of managing and maintaining complex oil and gas process systems, means that effective asset management can present significant challenges. For many businesses, partnering with an experienced, knowledgeable and specialized partner provides a far more cost-effective option. Outsourcing all or strategic parts of the process can deliver greater flexibility, accountability and control; it can also relieve the pressure on existing resources, such as freeing up internal engineering teams to concentrate on other business-critical activities. The benefits of an outsourced partnership. One such

organization that decided to outsource its asset management process with an experienced partner is a major US-based oil

FIG. 1. A detailed AEO plan addresses four key areas to improve plant reliability and asset utilization: maintenance strategy, work identification, work control and work execution.

30 MARCH 2016 | HydrocarbonProcessing.com

and gas pipeline transportation company. The business operates more than 12,000 mi of pipeline (FIG. 2), 150 main pumping stations and a number of key distribution terminals at railheads, ports and road hubs. Over a 10-yr period, however, the company has undergone several mergers and changes in ownership. Senior managers recognized that this led to a gradual loss of focus on machine reliability, with inconsistent practices and methods of operation across the pipeline and distribution network. They also understood that the company’s predictive maintenance strategy required a complete revaluation, but that the business lacked the necessary in-house skills and resources to carry this out effectively. The partner was commissioned to carry out a CNA study, and then to provide condition monitoring services using its extensive network of field-based service technicians to assess the status of 700 pipeline assets across North America. The partner began investigating the most valuable assets, primarily at a number of key distribution terminals, focusing on critical systems and equipment. An important element in this procedure was the use of its RCM techniques, which concentrated on dominant failure modes and the effects of these failures. Specific actions were recommended to prevent problems from reoccurring. Noncritical events were also evaluated and appropriate actions taken to allow the customer to optimize maintenance costs and increase productivity. The same approach is now being applied to the pipeline network of pumping stations. The partner subsequently worked with the customer to begin developing standard job plans that defined the critical steps that were required for each monitoring and maintenance activity: e.g., the repair of pump motors, including a list of the tools and parts needed, the repair steps and correct sequencing involved, and the time and resources required. This plan will be extended still further with a spare parts and stores optimization program (SPO), which minimizes stockholding and costs while improving the availability and location of key components to ensure that repairs are carried out quickly and cost-effectively. This strategic approach to asset management has had both short- and long-term benefits for the customer. Outsourcing the management of condition monitoring services—paid for

FIG. 2. A US oil and gas transportation company, which operates more than 12,000 mi of pipeline, turned to a partner because it understood that it lacked the in-house resources to effectively carry out a complete evaluation of its predictive maintenance strategy.

Project Management via an agreed monthly management fee—allowed the customer to move the costs from capital expenditure (CAPEX) to operating expenditure (OPEX) budgets. This makes it far easier to justify the cost of the program, while improving accountability proactivity and cash flow. The partner’s strategic approach to asset management provides a clearly defined and consistent operating methodology that can easily be adapted as the needs of the customer’s business or the operating environment change, providing a solution that is both secure and potentially future-proof. Perhaps most importantly, a savings of more than $1 MM was delivered in the first 12 months of the contract through improved asset uptime and productivity, and reduced repair and maintenance costs. This savings far outweighs the annual cost of the service contract. Experience, knowledge and resources. Across the indus-

try, outsourcing is becoming increasingly common for core business services ranging from facilities management to logistics and information technology (IT). For companies in the oil and gas sector, the challenge in outsourcing the management of mission- and safety-critical assets is to find a partner that has the requisite knowledge, experience and global resources to provide the reassurance that each and every asset will be monitored and maintained safely and to the highest standards at all times. These suppliers are able to provide this reassurance. Just as importantly, because they have specific and specialized skills in the field of asset management, condition monitoring and

preventive maintenance, they are able to deliver better results, faster and more efficiently than a comparable in-house function. An outsourced asset management partner adds an extra strategic dimension to the work of in-house maintenance and engineering teams, and can bring a fresh passion, greater strategic focus and a sense of purpose. Ultimately, the outcome should be improved operational efficiencies, cost savings and increased levels of asset optimization that allow oil and gas companies to become increasingly agile in response to growing market volatility. At a recent program-review meeting with the pipeline customer, a stakeholder applauded the fact that no asset under the monitoring program had experienced a failure in three years. 1

LITERATURE CITED PwC Strategy&, 2015 Industry Perspectives, http://www.strategyand.pwc.com/ perspectives/2015-oil-gas-trends. CHUCK NEWMISTER is an account manager for reliability end-user customers with SKF USA. He has two decades of experience in the oil and gas and hydrocarbon processing industries, and has been involved in machinery reliability sales and service centered on the US Gulf Coast for most of that time, serving with SKF for five years. Mr. Newmister is a member of the Houston Chapter of the Society of Maintenance and Reliability Professionals and has presented related material at numerous conferences. He earned a BS degree in physics at Central Methodist University in Fayette, Missouri.

EFFICIENCY MATTERS RESTORING PLANTS QUICKLY AND SAFELY Managing turnarounds on time and on budget can present many challenges. Cudd Energy Services helps you meet these challenges head on. Our fleet of pumping, transport, and storage vessels accommodates a wide range of flow rates for HPHT, open-flame environments that get you back online safely and efficiently. With a 1.5 million scf capacity, the queen storage vessel tank reduces frequent deliveries that cause congestion, and can be safely replenished without interrupting pumping operations, saving you time and money. Equipped with emergency shut-down devices, the dual mode pump features a heat recovery system that reduces fuel costs, and its EPA Tier 2/C.A.R.B. emission rating helps reduce emissions. For more information about our industrial nitrogen solutions, visit us at www.cudd.com or call us at 832.452.2800.

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Hydrocarbon Processing | MARCH 2016 31

Save 15% with Super Early Bird Registration when you Register by 22 March

6–8 June 2016 | Milan Marriott Hotel–Milan, Italy | HPIRPC.com

Discover Innovations in the Downstream Hear From Leading Operators Including: Total, Shell Global Solutions USA, Indian Oil Corporation, Petrobras + More We’re pleased to announce the agenda for Hydrocarbon Processing’s International Refining and Petrochemical Conference (IRPC), to be held 6–8 June in Milan, Italy at the Milan Marriot Hotel. The theme for this year’s conference is “Innovation in the Downstream.” The multi-track program features sessions devoted to: process technology / process optimization / plant design; internet of things; maintenance and reliability; efficiency and optimization; gas processing; catalysts; clean fuels / biofuels; heavy oil; refining / petrochemical integration; training/safety; water treatment / effluence management; feedstocks / alternative feedstocks; and emerging technologies. As an IRPC attendee, you will hear from key players in the global petrochemical and refinery sectors regarding the latest industry advancements, best practices, and case studies. In addition, you’ll have numerous opportunities to network with top operators and technology leaders from across the global hydrocarbon processing industry (HPI).

Comprehensive Program Features 3 Tracks & 60+ Presentations Participants include high-level professionals from these top operating companies + more: • Total • Eni • Shell Global Solutions USA • BayernOil

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• OMV • PDVSA • Indian Oil Corporation Limited • Petrobras

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• Oil & Natural Gas Corporation Ltd (ONGC) • SINOPEC

Exploring the Forefront of HPI Technology Developed by our Esteemed Advisory Board, the 2016 Agenda Features these Innovative Presentations: • Standard interface between TOTAL fuel blend optimizer ANAMEL and commercial oil movement automation suites — Edith di Crescenzo, Project Manager Advanced Control and Optimization, TOTAL • New reactor internals can be used to enhance profitability — Pankaj H. Desai, Licensing Sales Manager, Shell Global Solutions (US), Inc • Energy efficient way to debottleneck LP section of hydrocrackers with enhanced LPG recovery to improve profitability — Anurag Sharma, Deputy Manager – Process, Indian Oil Corporation Limited • Removal of nitrogen compounds from a FCC feed charge — Carlos Fernando Pinto Machado e Silva, Petrobras • EPRES technology of hydrogenation catalyst and Its commercial application — Yulan Gao, Fushun Research Institute of Petroleum and Petrochemicals, SINOPEC • Novel use of spent amine solution of gas sweetening plant as H2S scavenger, corrosion inhibitor & biocide — Dr. A K Shukla, Chief Chemist, Oil & Natural Gas Corporation Ltd (ONGC) • Eni slurry technology: Achieving the goal of residue conversion — Nicoletta Panariti, Licensing Manager, eni

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Petrochemicals

BEN DUBOSE, DIGITAL EDITOR [email protected]

Investigation into West Fertilizer blast shows room for safety improvement The final investigation report on the massive West Fertilizer Co. fire and explosion in 2013 serves notice to the chemical industry that, while progress has been made, significant gaps often still remain within site safety protocols. The fire and explosion on April 17, 2013, in West, Texas, resulted in 15 fatalities, more than 260 injuries and widespread community damage (FIG. 1). After investigating the incident for more than two years, the US Chemical Safety Board (CSB) released its final incident report in early February. Many of its findings were quite jarring, to say the least. What led to the explosion. The deadly incident occurred when about 30 tons of fertilizer-grade ammonium nitrate (FGAN) exploded after being heated by a fire at the storage and distribution facility. On the night of the explosion, the CSB found that there was a stockpile of 40 to 60 tons of ammonium nitrate stored at the facility in plywood bins. Though FGAN is stable under normal conditions, it can violently detonate when exposed to contaminants in a fire. “This tragic accident should not have happened,” said Vanessa Allen Sutherland, CSB chairperson. “We hope that by sharing lessons learned from our West Fertilizer investigation, we will help raise awareness of the hazards of fertilizer grade ammonium nitrate.” The CSB is an independent US federal agency charged with investigating serious chemical accidents. While the board does not issue citations or fines, it regularly makes safety recommendations to companies, industry organizations, labor groups and regulatory agencies. Lack of community awareness. The

CSB also found that several factors contributed to the severity of the explosion, in-

FIG. 1. The nearby community in West, Texas, suffered extensive damage after the April 2013 explosion at the West Fertilizer Co. plant.

cluding poor hazard awareness and the fact that nearby homes and businesses were built in close proximity to the plant over the years prior to the accident. “We found that, as the city of West crept closer and closer to the facility, the surrounding community was not made aware of the serious explosion hazard in their midst,” said Johnnie Banks, team lead for the CSB’s investigators. “West Fertilizer Co. underestimated the danger of storing fertilizer-grade ammonium nitrate in ordinary combustible structures.” Investigators concluded that this lack of awareness was due to several factors, including gaps in federal regulatory coverage of ammonium nitrate storage facilities. Subpar emergency planning. Finally,

the CSB also said that inadequate emergency planning contributed to the tragic accident. Investigators found that the West Volunteer Fire Department was not required to perform pre-incident planning for an ammonium-nitrate-related emergency, nor were the volunteer firefighters required to attend training on responding to fires involving hazardous chemicals.

As a result, the CSB made several safety recommendations to various stakeholders, including the US Environmental Protection Agency (EPA), to better inform and train emergency responders on the hazards. More specifics in video. The specific

safety recommendations and proposed policy changes by the CSB are available in a 12-minute video posted on the HPInformer blog at HydrocarbonProcessing.com. The video, entitled “Dangerously Close: Explosion in West, Texas,” includes a 3D animation of the fire and explosion, as well as interviews with CSB investigators and chairperson Sutherland. While the damage in West has already been done, the CSB’s hope is that the publicity from this investigation and its findings can reduce the potential for similar incidents in the future. “The CSB’s goal is to ensure that no one else be killed or injured due to a lack of awareness of hazardous chemicals in their communities,” Ms. Sutherland said. “If adopted, the board’s recommendations can help prevent disasters like the one in West, Texas.” Hydrocarbon Processing | MARCH 2016 35

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Engineering Case Histories

TONY SOFRONAS, CONSULTING ENGINEER http://mechanicalengineeringhelp.com

Case 89: Cracking of welds due to weld fatigue Pressure vessels, piping and fabricated structures adhering to recognized welding codes rarely fail in static loading, unless they are damaged or severely overloaded. Weld cracking can develop in pressure vessels under cyclic conditions.1 While extensive research and design guides exist for the design of welds under cyclic conditions, the complex metallurgy of a weld, and the quality of the weld, are difficult to define. Here, in-service failures are discussed, based on the author’s experience with fatigue cracking of welded fabrications. Calculations2 had been performed to determine the nominal cyclic stress present when failures occurred. In this manner, limiting cyclic stress could be determined and compared with available test data. Why welds fail in fatigue. Metallurgical

changes, shrinkage, residual stresses, stress concentrations and internal defects can combine to cause a lower endurance limit in fatigue from the base metal. Fatigue can be thought of as a tensile stress that opens up a preexisting crack that may be small, causing it to grow. Here, a plus symbol (+) is used to show a stress cycling from zero to a tensile stress, indicating the opening of a crack. Weld fatigue life. The author has seen many in-service fatigue failures. FIG. 1 illustrates one of many growing cracks in the toes of fillet welds on a vibrating conveyor that had undergone misalignment and twisting. The failure occurred with only +6,000 lb/in.2 on several marginal welds within one year’s operation. Eliminating the misalignment dropped the stress to +1,500 lb/in.2, and the conveyor operated for 10 years without further cracking. FIG. 2 shows a failure in a stainless steel pipe shaft undergoing a rotating bending stress. Notice how the weld starts from one side at a defect and progresses through the shaft. This particular weld had a small initial crack that started to propagate in less

FIG. 1. Crack in toe of fillet weld.

FIG. 2. Fatigue weld pipe defect.

than 1,000 cycles at +30,000 lb/in.2 There is good reason to be suspicious whenever an impact load causes a nominal stress of +30,000 lb/in.2 or higher on a weld. FIG. 3 is a bending failure in a plug-type weld, which starts at an existing design fabrication crack in the “blind” zone that was inaccessible. The weld melt outline is visible and was blended flat on the left plate. Notice that there is a gap that acts as the initiation point for the crack. Due to this design issue, the nominal bending stress on the plate must be kept low. This was achieved by using a thicker plate and better weld geometry. All of these failures had a calculated nominal cyclic stress of more than +6,000 lb/in.2 Once a fatigue crack starts, it usually continues to grow at these stress levels. How to reduce fatigue failures. Designing and operating procedures to keep cyclic weld stresses below those cited are helpful. For example, on vibrating conveyor specifications, an equipment manufacturer agreed that, for critical welds, the nominal stress should be less than +3,000 lb/in.2, as verified with strain gauges. On welded augur shafts, cleaning of buildup

FIG. 3. Blind plug weld.

dropped cyclic stresses from +8,000 lb/ in.2 to zero. Other techniques are available for reducing fatigue failures.3 LITERATURE CITED Sofronas, A., B. Fitzgerald and E. Harding, “The effects of manufacturing tolerances on pressure vessels in high cycle service,” ASME, Pressure Vessels and Piping, Vol. 347, July 1997. 2 Sofronas, A., Analytical Troubleshooting of Process Machinery and Pressure Vessels, Including Real-World Case Studies, John Wiley & Sons, January 2006. 3 Maddox, S. J., “Fatigue strength of welded structures,” Woodhead Publishing, January 1991. NOTE Case 88 was published in HP in January. For past cases, please visit HydrocarbonProcessing.com. 1

TONY SOFRONAS, D. Eng, was the worldwide lead mechanical engineer for ExxonMobil Chemicals before retiring. He now owns Engineered Products, which provides consulting and engineering seminars on machinery and pressure vessels. Dr. Sofronas has authored two engineering books and numerous technical articles on analytical methods. Hydrocarbon Processing | MARCH 2016 37

| Special Report CORROSION CONTROL Corrosion is often the root cause of equipment failure, process unit and plant downtime, off-spec product generation, contamination and accidents. In the hydrocarbon processing industry (HPI), corrosion is an ongoing and dynamic issue for processing facilities and equipment. Corrosion-related damage is accelerated by several factors, such as high temperatures, acidic/caustic conditions and erosive fluids; all are found in HPI facilities. Likewise, aging process equipment is vulnerable to corrosion attacks unless preventive and maintenance measures are applied on a regular basis. The special report investigates methods to mitigate corrosion attacks on HPI processing equipment and infrastructure, along with the use of new and updated technologies to extend equipment life and prevent corrosion. Photo courtesy of Valero Energy Corp.

Special Report

Corrosion Control D. LEE, J. KLINKENBIJL and T. BROK, Shell Global Solutions, Amsterdam, The Netherlands; and J. CRITCHFIELD and D. VALENZUELA, Shell Global Solutions, Houston, Texas

Improved corrosion prevention with acid-aided regeneration technology A commonly applied process for the removal of acid gas contaminants from gas is amine treating. At its heart lie the absorption of acid gases into an amine solution and the regeneration of this solvent to be fed back to the absorption column. The interplay between these two steps is important, as the treating performance can be majorly influenced by the amount of acid gas dissolved in the lean solvent returning from the regenerator. In some cases, achieving the treated gas specification using standard reboiler duties in the regenerator column is difficult, as the loading of the solvent entering the absorber may not be lean enough. To obtain a leaner solvent, the regeneration can be enhanced by the implementation of proprietary tail gas treating technology that removes sulfur compounds downstream of the Claus unit. In other situations, this technology offers the ability to optimize amine processes by achieving a certain solvent leanness with lower regeneration steam requirements. The proprietary technology relies on applying a controlled concentration of certain acidic additives that enhance regeneration. Elements of the technology were originally developed more than 40 years ago; the technology has evolved in application over the past decades to incorporate corrosion-avoidance and improved-operability strategies. The mature technology is particularly useful in existing plants, as it can often help meet tighter specifications without requiring hardware changes, thereby offering valuable benefits in a world in which specifications and emissions limits change over time, and the pressure to reduce capital expenditures (CAPEX) and operating expenditures (OPEX) is ever increasing. Here, the application of the technology in actual operation is examined, and trends and observations are drawn from operating experience.

PRINCIPLES In absorption, the acid gas reacts with the amine in an equilibrium reaction to form ions; for example, H2S is removed by reaction with a tertiary amine, as shown in Eq. 1: H2S(g) D H2S(aq) + R3N D HS– + R3NH+

(1)

This reaction is then reversed in the regenerator, stripping the acid gas out of the solvent. In the bottom of the column, this reverse reaction is favored by a lower pH, and, therefore,

by the addition of acidic regeneration-enhancing agents—for example, a strong mineral acid, as shown in Eq. 2: HX + R3N D X– + R3NH+

(2)

The R3NH+ concentration increases, forcing the equilibrium of the reaction in Eq. 1 to the left and leading to a greater release of H2S—and, therefore, a leaner solvent leaving the regenerator.

EXAMPLES The technology company’s experience with adding regeneration-enhancing additives to amine treating systems is mainly based on its tail gas treating units (TGTUs). These units remove H2S, using a selective amine solution to recycle it back to the sulfur recovery unit, thereby minimizing SO2 emissions from the plant. In the process, deep solvent stripping is applied to remain comfortably within environmental limits. Good plant data describing plant performance and solvent analysis are required to successfully apply acid-aided regeneration. A comparison of the experience at several different amine units clearly showed the effect of additive presence, and the observations from these units and case studies are presented here. Case studies 1 through 6 address plants removing H2S (with a focus on TGTUs, but also discussing high-pressure applications). Case Study 7 reviews the technology company’s experience of applying regeneration enhancement to CO2-only applications. Case Study 1: Lean H2S loading. This case study concen-

trates on improving the controllability of H2S lean loading in multiple TGTUs. The unit’s regenerator performance was optimized at various locations to decrease SO2 emissions. Regeneration enhancement was applied in the multiple TGTUs, which resulted in improved H2S performance of the absorbers. Structured performance demonstrations were also conducted; the concentration of regeneration-enhancing additive was manipulated in controlled tests. H2S performance in the absorber (and, therefore, SO2 performance of the overall tail gas treating) was improved in each system where the technology was implemented. From this experience, an overall relationship was established between the leanness of the solvent and the H2S performance of the TGTU absorber. However, the operability of the overall TGTU system was also considered in the demonstrations. Hydrocarbon Processing | MARCH 2016 39

Corrosion Control It was discovered that when the additive concentration increased beyond a certain maximum value, regenerator performance became increasingly sensitive to variation in applied steamrate. This increased sensitivity to energy input is demonstrated conceptually in FIG. 1, which shows trends observed from the collated plant test data from the TGTUs. FIG. 1 illustrates the overall relationship between the amount of regeneration-enhancing additive, energy input and resulting H2S lean loading. It was observed that, with the same energy input, the solvent becomes leaner when a higher concentration of additive is employed; and, at the same leanness, less steam is required with higher additive concentration. Case Study 2: Absorber performance. Enhancing regeneration in an amine unit can benefit treating performance only if the absorber operates at a close approach to lean solvent loading. In practice, this means that the absorber must have enough stages to allow for deep treating of H2S. An additional requirement is that improvement in regenerator performance must be greater than the loss in the absorber top. If these requirements are not met, then performance in the absorber can actually worsen when acidic additives are present. Lower additive concentration

H2S in lean solvent

H2S too insensitive to steamrate

Higher additive concentration H2S too sensitive to steamrate

This is because the presence of the additives also results in a higher H2S vapor pressure, which disfavors absorption. The strength of this impact depends on several interacting factors, such as the relative amounts of H2S, acidic additive and CO2 in solution, along with the design and operating conditions. In TGTU absorbers, the condition of the top trays is most important for control of absorber performance, and, in this region, coabsorption of CO2 plays a large role. CO2 interferes in H2S removal, thereby decreasing the effect of enhanced regeneration. If substantial amounts of CO2 are coabsorbed, then the approach to H2S equilibrium over the lean solvent worsens, minimizing the deep H2S treating benefit of the technology. This effect is illustrated in FIG. 2 from simulation, which is compared against known data;1, 2 partial pressure of H2S increases with increasing CO2 coabsorption. Case Study 3: TGTU with near-flooding conditions. After a major capital change in a TGTU, the amine regenerator began to experience near-flooding conditions, which posed difficulty in achieving deep H2S stripping. In this particular situation, the operators found that changing the steamrate had only a minor effect on the depth of H2S stripping. When the ambient temperature was high, this unit experienced challenges meeting environmental targets. A project to address the limitations concluded that the best remedy was to formulate the methyldiethanolamine (MDEA) solvent with the appropriate dosing of proprietary additive. This action succeeded in reducing the regeneration steam by 40% to achieve the desired depth of stripping while the specification of the treated gas remained unchanged, which also lowered the regenerator differential pressure and reduced the risk of flooding (FIG. 3). No change was observed in the decomposition rate of the solvent, and no noticeable accumulation of soluble materials was found. Also, following formulation, the operators now observe a clear relationship between the steamrate and the resulting H2S concentration in the treated gas, improving operability of the unit.3 Case Study 4: Energy savings. Steam savings possible for an example case are illustrated in FIG. 4. Based on the previously outlined conclusion to target additive content for controllabil-

Steamrate

FIG. 1. Summary of TGTU test data results.

0.0

0.3

Relative steam ratio Relative stripper differential pressure Relative H2S in treated gas, ppmv

0.0

40°C 60°C GPA data, 40°C

0.0

H2S partial pressure

0.2 0.0

Formulation period

0.0

0.1

0.0

CO2 loading

FIG. 2. Effect of CO2 coabsorption on H2S absorption.

40 MARCH 2016 | HydrocarbonProcessing.com

FIG. 3. Effect of formulation on differential pressure, steamrate and H2S performance.

30-May-08

28-May-08

1.0

26-May-08

0.8

24-May-08

0.6

22-May-08

0.4

20-May-08

0.2

18-May-08

0.0

16-May-08

0.0

14-May-08

0.0

Corrosion Control

ence is with TGTUs, it has also applied enhanced regeneration in other applications, including selective H2S removal in highpressure natural gas. As with the tail gas treating technology, that application requires deeply regenerated lean solvent to meet the H2S specification in the treated gas. This effect is illustrated in Case 5, in which H2S performance of the system was optimized by manipulating the extent of regeneration enhancement. FIG. 5 shows the result of changing the amount of the additive: low H2S lean loadings that occur at high additive concentrations on an example high-pressure integrated system. This plant was dependent on regeneration enhancement to reach targeted H2S concentration in the lean solvent. Test runs demonstrated that higher levels of enhancement resulted in too-deep regeneration—deeper than the company’s technical governance for the application. Case Study 6: Inferring corrosion effects. It is well known

Energy savings/steam ratio

that amine systems processing H2S tend to form a layer of iron sulfide on exposed carbon steel (CS) surfaces. This layer is thought to help protect CS surfaces against certain types of corrosion. Past publications have introduced the concept that maintaining a minimum level of H2S in the lean solvent facilitates preserving this iron sulfide layer.4, 5 This operating philosophy was applied to the plant in Case 5. The depth of regeneration was maintained on target by controlling the amount of regeneration enhancement. Monitoring of corrosion coupons, filter changes and solvent-quality tests demonstrated a substantial decrease in corrosion in that location. Building on that success, a systematic review of tail gas treating corrosion experience was conducted in the company’s US downstream applications to document CS corrosion in key locations within the amine units. This review was coupled with solvent quality monitoring in the locations. The study demonstrated that plants maintaining H2S in the solvent were less likely to detect iron in solvent quality analysis.

Additional steam savings sacrificed for improved operability 16% of original

Steam savings in example case TGTU with AAR 40% of original Required steam ratio Steam savings Acid concentration

FIG. 4. Steam ratio savings.

Case Study 7: Regeneration enhancement in CO2removal systems. The technology company has also reviewed

the effect of regeneration-enhancing additives with the objective of reducing the regeneration heat requirements, both in addition and in comparison to other options to reduce the energy footprint of the unit at a specific plant. Although the CO2 speci-

H2S in lean solvent

Case Study 5: Application in high-pressure selective treating. Although most of the technology company’s experi-

Testing for iron in the solvent samples can give an indication of possible corrosion. Although corrosion may occur without the solubilization of iron, if iron is found in the solvent samples in H2S removal plants, this can be seen as a warning sign of conditions that may lead to corrosion. Plant data shows that iron content correlates with high acid concentration, as well as low concentration of suppressive H2S in the sample. FIG. 6 shows the relationship between iron content, depth of stripping and additive concentration in the population of TGTUs. A simple relationship is observed: plants that do not strip too deeply are less likely to find iron in solvent samples.

Performance target range

Regenerator enhancement

FIG. 5. Relationship between H2S lean loading and additive concentration in a high-pressure application.

Iron not found in solvent samples

H2S leanness

ity, the bulk of the possible steam savings can still be realized, as the steam savings curve flattens at higher acid concentrations. In this example case, 40% of the steam can be saved while maintaining improved operability and preserving a minimum concentration of H2S in the lean solvent.

Iron occasionally in solvent samples

Iron found in solvent samples

Regeneration enhancement FIG. 6. Iron prevalence in MDEA-based TGTUs. Hydrocarbon Processing | MARCH 2016 41

Corrosion Control

Acid gas lean loading, mol/mol

H2S AAR DIPA CO2 AAR DIPA H2S normal curve CO2 normal curve

Steamrate

FIG. 7. H2S and CO2 lean loading in DIPA before and after adding additives.

CO2 leanness, mol/mol

Low acid content Medium acid content High acid content

in a structured way. The data suggests that the presence of acids in the solvent significantly impacts CO2 lean loading to some extent. This data presents another example of the effect of acid content in high-pressure absorption, although it is not as pronounced as in the case of H2S, for the reasons outlined above.

TAKEAWAY Enhancing regeneration in amine treating systems has proven beneficial in different applications by improving operations and relieving design limitations through reduced steamrates and/or improved treating performance. However, care must also be taken, since an improper dosing of acid can lead to corrosion risk, worse treating performance and reduced controllability of the unit. From experience in operating with acidic additives and controlled plant tests, several observations were reviewed to understand how to avoid corrosion and improve treating results while applying regeneration enhancement. NOTE The tail gas treating technology referenced in this article is Shell Claus Offgas Treating (SCOT) technology, and the proprietary units referenced are Shell SCOT units. LITERATURE CITED Huang, S. H. and H. J. Ng, “Solubility of H2S and CO2 in alkanolamines,” GPA Research Report RR-155, September 1998. 2 Bullin, J. A., R. R. Davison and W. J. Rogers, “The collection of VLE data for acid gas—alkanolamine systems using Fourier transform infrared spectroscopy,” GPA Research Report RR-165, March 1997. 3 Bonner, S. and J. Critchfield, “Relieving stripper flooding at Martinez SCOT 3,” presented at Brimstone Sulfur Symposium 2009. 4 Van Roij, J., J. Klinkenbijl, P. Nellen and K. Sourisseau, “Materials threats in aging amine units,” Paper 2207 presented at NACE Corrosion 2013. 5 API Recommended Practice 945, “Avoiding environmental cracking in amine units,” April 2008. 1

Steamrate

FIG. 8. Acid impact on CO2 solvent leanness and steamrate on an example unit.

fications are somewhat more relaxed, the actual performance examines a similar deep removal compared to the usual H2S specification partial pressures. Plant test data in a secondary amine solvent system was taken and is presented in FIG. 7, which predicts that, in conditions similar to TGTUs, steam savings exist for CO2 , but are much lower than for H2S and follow a much flatter curve. With the lower steam savings, adding acids looks less interesting for CO2 systems—particularly since CO2 is generally easier to strip out than H2S, and since CO2 specifications are often less severe and do not require deeper stripping of CO2 . Note that a secondary amine forms carbamates in the presence of CO2 , a relatively more stable component, which is more difficult to regenerate and makes it difficult to strip to very low CO2 . FIG. 7 illustrates the effect of the additives. The lines indicate trends derived from plant data, while the performance with additives is shown in data points from the tests. The data indicate that systems using a secondary amine such as DEA or DIPA can also benefit from the addition of acids, but less data are available than for the tertiary MDEA solvents. Observations have also been made for an MDEA-based solvent and are presented in FIG. 8. In this example, a gas treating unit processing high-pressure gas for deep CO2 removal was analyzed, based on plant performance vs. plant leanness analysis, 42 MARCH 2016 | HydrocarbonProcessing.com

DANMI LEE joined the gas processing group at Shell in 2014. Upon completing her master’s degree in chemical engineering at University College London in the UK, she started working for an engineering firm, performing concept studies and preliminary designs for cryogenic gas processing technology. At present, she is working on the development and validation of amine treating models within Shell’s gas processing technology team. JEANINE KLINKENBIJL joined Shell’s gas treating group in 1982. Her expertise includes absorption, adsorption and sulfur conversion, as well as carbon capture and storage, covering the operation, design and development of new processes, lineups and equipment. Presently, she is team lead in gas processing expertise and also holds the gas treating principal technical expert role at Shell. Some of her previous roles include positions in process modeling, crude characterization, refinery support and IT. She holds a master’s degree in chemical technology from the Technical University in Eindhoven, The Netherlands. THEO BROK is a senior deployment engineer in the gas processing group at Shell in The Netherlands. He is also a subject matter expert for amine and caustic processes. He is leading the development of the X-solvents (ADIP-X and Sulfinol-X) and has wide experience in sour gas projects. Mr. Brok holds a master’s degree in chemical engineering from the Technical University of Eindhoven in The Netherlands. He has worked for Shell for 27 years, including eight years in LNG and sour gas processing facilities in Brunei and The Netherlands. JIM CRITCHFIELD is a senior process engineer at Shell Global Solutions in Texas. DIEGO VALENZUELA is a principal process engineer in gas treating and sulfur at Shell Global Solutions in Texas.

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Regional report

M. RHODES, Managing Editor, and M. NOGARIN, Contributing Writer

Central American nations beef up import infrastructure, fuel production amid demand shift Due to the growth in the region’s middle class, Central America and Mexico have seen tremendous petroleum product demand growth over the past decade. Multiple forecasts show that the region will see a nominal increase in demand through the rest of the decade. Demand has been shifting to more middle and light distillates, as opposed to fuel oil. Unfortunately, the region’s refineries have not been able to keep up with demand and are challenged to produce higher-grade, lower-sulfur transportation fuels. Coupled with the shortage of fuel production, the drop in oil prices has hammered the region, especially those countries that depend heavily on oil export revenues. To satisfy the growth in demand, Central America and Mexico have relied heavily on refined fuel imports from the US, as the drop in crude oil prices has left little money to fund

capacity expansions. In the short term, these nations would rather import refined fuels than invest in major expansions or grassroot facilities, which can be multibillion-dollar endeavors. This trend does not mean that the region is devoid of refining projects. A forecast slowdown in regional demand over the next few years will reduce refined fuel imports; however, should regional growth pick up, major consumers will need to increase refined fuel imports due to inadequate refining capacity. This will include the expansion and upgrade of multiple refineries; the construction of methanol, ammonia-urea and petrochemical plants; and the development of new LNG infrastructure.

MEXICO Mexico is the tenth-largest oil-producing country in the world, with approximately 2.8 MMbpd of output.

However, the country has witnessed production declines over the past decade due to natural declines in mature fields. According to BP’s Statistical Review of World Energy 2015, Mexico’s oil production has declined from approximately 3.8 MMbpd in 2004 to 2.8 MMbpd in 2014. To combat oil production declines, the country has instituted energy reforms that could be the first step in a deeper involvement among international firms in Mexico. These reforms mark a radical shift away from state-owned Pemex’s 75year monopoly on the oil and gas market, potentially opening up some of the world’s biggest remaining untapped oil reserves to private companies. The energy reform plan would see more than 900 onshore and offshore fields auctioned off over the next five years. Mexico hopes this will attract up to $63 B in investment by 2018 and increase domestic production by 1 MMbpd by 2025.

Left: Methanex’s methanol facility in Trinidad. The company operates two methanol plants on the Point Lisas Industrial Estate on Trinidad’s west coast. Photo courtesy of Methanex. Right: Construction of the Etileno XXI project in Mexico. The plant is scheduled to begin operations by the end of 1Q 2016. Hydrocarbon Processing | MARCH 2016 45

Regional Report The country’s first auction, held in July 2015 and primarily consisting of shallow-water fields, was snubbed by supermajors. Major oil-producing companies like ExxonMobil, Chevron and Total are opting for deepwater fields with substantial proven oil reserves. The auctions for these fields began in late 2015. Should a production boom occur in Mexico—which, according to recent estimates, also has the world’s fourth-biggest shale gas reserves—further downstream development could occur by the start of the 2020s. Local producer Mexichem has already said it is targeting numerous ventures and acquisitions with foreign companies, although weaker oil prices may slow those developments as companies reevaluate the market. Refining. Pemex owns and operates six

refineries in Mexico with a combined installed capacity of approximately 1.7 MMbpd. These facilities process light and extra-light crudes produced domestically. Heavier domestic crudes are exported, mainly to the US Gulf Coast (USGC) for processing. Domestic output is sufficient to meet consumption, but the country lacks adequate refining capacity to satisfy demand for transportation fuels. As a result, Mexico is forced to import refined petroleum products. Imports of gasoline averaged nearly 400 Mbpd in the first half of 2015. Illegal taps on gasoline pipelines forced Mexico to boost gasoline imports by an additional 75 Mbpd in mid-2015. The country had announced major refining capacity expansion investments;

however, due to the drop in oil prices, those plans have been delayed. Low oil prices forced Pemex to slash $4 B from its 2015 budget, including a much-needed program to expand and upgrade the country’s refining network. However, in late 2015, Pemex reinstated its domestic refining modernization program. The company plans to invest as much as $23 B to upgrade its refinery system to produce more clean fuels and to expand processing capacity. Upgrades at Pemex’s Tula refinery are already underway, with additional plans to upgrade the Salamanca and Salina Cruz refineries. Additional investments will be made to more than double the output of ultra-low-sulfur gasoline and diesel, as well as in cogeneration projects to boost electricity generation. The modernization program will rely on Pemex’s goal of seeking out private investors to fund refinery upgrades in exchange for a share of profits. Pemex ended 2015 with a total debt that exceeded $100 B. The company will need to rely on outside funding for these projects to move forward. Petrochemicals. Mexico has constructed its first major private-sector petrochemical project in 20 years (FIG. 1). Etileno (Ethylene) XXI is a $5.2-B greenfield petrochemical complex being built in Nanchital near Coatzacoalcos, Veracruz. The project consists of a 1-MMtpy ethylene cracker; two highdensity polyethylene (HDPE) plants, one with a capacity of 400 Mtpy and the other with 350 Mtpy of capacity, that

FIG. 1. Aerial view of the Etileno XXI project. Photo courtesy of Braskem Idesa.

46 MARCH 2016 | HydrocarbonProcessing.com

will also produce a full range of monomodal and bimodal high-density and medium-density polyethylene resins; one low-density polyethylene (LDPE) plant that will produce 300 Mtpy; and storage, waste treatment and utility facilities. The project also includes a 150MW combined-cycle power (CCP) and steam cogeneration plant; a multimodal logistics platform for the shipment of 1 MMtpy of polyethylene via bulk train or truck; and administrative, maintenance, control room and ancillary buildings. The JV company, Braskem Idesa has built, developed and will operate the production facility. Brazil-based Braskem is the largest producer of thermoplastic resins in the Americas, and Idesa is a leading Mexican petrochemical company. The project is intended to increase Mexico’s domestic petrochemical production to satisfy demand and reduce imports of petrochemical products. A glaring gap exists between Mexico’s investment potential for polyethylene production and its inability to meet surging demand. At present, 65% of polyethylene demand is satisfied through imports, and the gap continues to grow each year. The Etileno XXI project is forecast to replace $2 B of polyethylene imports used as feedstock for the agricultural, automotive, construction and consumer industries. The facility is the largest project finance transaction in the history of the petrochemical industry in Central America, as well as the biggest foreign investment in Mexico by a private Brazilian company. Etileno XXI is expected to begin operations by the end of 1Q 2016. Additional petrochemical projects include Pemex’s plan to restart the country’s fertilizer industry, which was nearly wiped out due to high natural gas prices. Pemex’s petrochemical division, Pemex Gas y Petroquímica Básica (PGPB), is looking to partner with fertilizer producers to increase the country’s production of ammonia and urea. PGPB is investing more than $230 MM in upgrades at Ferquimex’s 132-Mtpy Camargo ammonia plant in Chihuahua and in Cobra Instalaciones Mexico’s 1.5-Mtpd urea plant in Coatzacoalcos. Both projects are expected to be completed in early 2016. These two projects are a step in the right direction, but much more time and investment will be needed to restart the country’s fertilizer industry.

Regional Report Four cooling towers have a total of 32 cells: 15 cells for ammonia plants 4 and 5, and 17 cells for plants 6 and 7. Two demineralizer units are capable of producing 15,262 cmd of steam for the ammonia plants. Two electrical turbogenerators provide 29.8 MW of power each. The complex’s 6.6-Mt storage system consists of six spherical tanks, four with a working storage capacity of 1.2 Mt and two with a 900-t storage capacity. The

facility has the ability to send ammonia by pipeline to the refrigerated terminals at Pajaritos and Salina Cruz, thereby enabling ammonia output to either the Gulf of Mexico or the Pacific Ocean. The petrochemical complex also has two filling stations for either ammonia tank trucks or tank cars (rail). Urea production for the petrochemical complex comprises the following components: urea trains 1 and 2, which 19

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Mexico consumes 4 MMtpy of fertilizers. The country was self-sufficient in its production of fertilizers until the end of the 1990s; however, with increasing natural gas prices and the closing of numerous fertilizer-producing plants, Mexico was forced to become a net importer. Ninety percent of the country’s 1.5 MMtpy of urea consumption—a major component in the production of fertilizers—is imported. Taking into account that Mexico has 54.9 MM hectares of agricultural land under production, the country is targeting the expansion of its ammonia and urea markets. In 2013, Pemex hoped to increase the value-added price of natural gas with its purchase of fertilizer manufacturer AgroNitroginados. In August 2015, the administrative council of Pemex approved the activation of its subsidiary, Pemex Fertilizers. This made the construction of the Cosoleacaque petrochemical complex the main component of company’s fertilizer industry. The main objective of the Cosoleacaque plant has always been to expand and revitalize the fertilizers production industry in Mexico, so the upgrade project consists of rehabilitating the Urea production trains, the auxiliary services and the area of compression. The $220-MM plant modernization project began in September 2014, and it is estimated to be completed in March. Once completed, the facility will be able to produce 1 MMtpy of urea. The Cosoleacaque complex consists of the following components: ammonia plants 4, 5, 6 and 7, with production capacities of 1,440 tpd each. When the refurbishment project is completed this year, the complex’s ammonia production capacity will increase from 1.4 MMtpy to 1.9 MMtpy. The hydrogen (H2 ) recovery unit is fed with the purge gas from the ammonia plants. Ammonia is recovered and sent, in turn, to the storage tanks, and H2 and fuel gas are consumed again in the production units, while the absorption tower keeps ammonia from the relay processing plants to prevent emissions to the environment. Two storage tanks with a raw water capacity of 200 Mbbl and a 200Mbbl clarified water tank are part of the compound. For pretreatment, three units produce 130,824 cmd of clarified water for the cooling towers. The demineralizer units are also utilized as service water.

1

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3089_Catalysts_120x190_en.indd 1

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03.06.14  15:12

Regional Report each feature a production capacity of 1.5 Mtpd and transport to storage; a nitric acid plant that can produce 625 tpd; an

expand its 5,500-mi natural gas pipeline system, focusing on central and northern industrial cities. The majority of this

The drop in oil prices has left little capital to fund capacity expansions in Central America and Mexico. To satisfy demand growth, the region has relied heavily on refined fuel imports from the US. ammonium nitrate plant with a capacity of 818 tpd; and a plant for nitrogenated solutions, such as urea ammonium nitrate. These three plants are not included in this rehabilitation project, although their assets have been acquired. Many more plants will need to be refurbished and rebuilt if Mexico hopes to completely bring its fertilizer industry back to life. If the investments are made, then rebuilding this industry could take 5–10 years. Gas processing/LNG. Mexico relies

on natural gas imports from the US to satisfy domestic demand, mainly for power generation. According to the US Energy Information Administration (EIA), US gas exports to Mexico via pipeline have tripled in the past decade, hitting over 700 Bcf in 2014 (FIG. 2). Mexico suffered years of gas shortages as US pipeline capacity failed to keep up with its growing industrial demand for gas, and as Pemex focused on more profitable oil production. To satisfy increasing gas requirements, Mexico is investing more than $10 B to

network will be filled by natural gas production from US shale plays, primarily in southern Texas and in the Eagle Ford shale play. The country’s pipeline expansion projects have prompted Mexico to abandon plans to build nuclear power stations and to instead construct new CCP stations. These plans will provide additional power generation to meet the country’s increasing demand for electricity. With additional natural gas supplies pouring into the country, Mexico has announced plans to possibly export excess natural gas as LNG. Mexico now operates three LNG import terminals located at Altamira, Costa Azul and Manzanillo. Total domestic LNG regasification capacity is just over 15 MMtpy. Pemex is conducting feasibility studies on two projects that could turn Mexico into an LNG exporter. The first export terminal would be located near Salina Cruz on the Pacific coastline of Oaxaca. Feasibility studies on the proposed $6-B terminal began at the end of 2014, and the total liquefaction capacity is expected to be announced once the studies are completed. If built, the ter-

US natural gas pipeline exports to Mexico, MMcf

1,000,000

750,000

500,000

250,000

0 1980

1990

2000

2010

FIG. 2. US natural gas pipeline exports to Mexico have tripled in the past decade. Source: US EIA.

48 MARCH 2016 | HydrocarbonProcessing.com

minal could begin operations before the end of the decade. The second proposed project is to add liquefaction capabilities at Mexico’s existing Costa Azul LNG facility. Pemex has signed a memorandum of understanding (MoU) to develop the project with Sempra Energy units IEnova and Sempra LNG. Feasibility studies are taking place to examine the potential addition of more than 3 MMtpy of liquefaction capacity to the existing terminal.

TRINIDAD AND TOBAGO The island nation is the largest oil and natural gas producer in the Caribbean, and nearly half of the country’s GDP is tied to the energy sector. The country’s largest oil producer is state-owned Petroleum Co. of Trinidad and Tobago Ltd. (Petrotrin). Petrotrin also operates the country’s only refinery, the 168-Mbpd Pointe-à-Pierre facility, which is located on the west coast of Trinidad, approximately 56 km north of San Fernando. It produces liquefied petroleum gas (LPG), jet fuel, gasoline, diesel and fuel oil. Clean fuels projects. Construction was

recently completed on a new ultra-lowsulfur diesel (ULSD) unit at the refinery. The ULSD plant is part of Petrotrin’s clean fuels upgrade program to improve the profitability of the Pointe-à-Pierre refinery, as well as to meet new diesel quality specifications. Additional projects of the clean fuels program included a liquid fuel pipeline project and a gasoline optimization program. On the petrochemical side, Trinidad and Tobago is the world’s largest exporter of ammonia and the secondlargest exporter of methanol. The country has 11 ammonia plants and seven methanol plants.

Petrochemical output. According to Trinidad and Tobago’s Ministry of Energy, overall production and export for ammonia, methanol and urea totaled over 420 Mtpy in 2013. The country is investing $1 B in the construction of a new methanol and dimethyl ether (DME) production complex. The project is being developed by state-owned National Gas Co. of Trinidad and Tobago, Massy Holdings and a consortium consisting of Mitsubishi Gas

Regional Report Chemical, Mitsubishi Corp. and Mitsubishi Heavy Industries. The complex will be owned by Caribbean Gas Chemical, a JV of the aforementioned companies. The facility will be located in La Brea and have a total capacity of 1 MMtpy of methanol and 20 Mtpy of DME. The plant is expected to begin operations in 4Q 2018. Natural gas exports. Since the early

1990s, Trinidad and Tobago’s hydrocarbon sector has shifted from an oildominated sector to mostly natural gas. The country boasts one of the largest natural gas processing facilities in the Western Hemisphere. The Phoenix Park Gas Processors Ltd. natural gas liquids (NGL) complex has a processing capacity of nearly 2 Bcfd and an output capacity of 70 Mbpd of NGL. The products are transferred to various power plants for electricity production and to petrochemical plants for feedstock. Natural gas is utilized in many sectors of the country, including the production of LNG, feedstock for petrochemical manufacturing and metals refining, and the production of nearly all of the country’s electricity generation. The country also converts natural gas into LNG for export. At present, Trinidad and Tobago is the sixth-largest LNG exporter in the world. The nation exports more than 14 MMtpy of LNG from Atlantic LNG’s Point Fortin terminal. The plant consists of four liquefaction trains with a total installed liquefaction capacity of nearly 15 MMtpy. The trains vary in size from 3 MMtpy to more than 5 MMtpy. LNG exports are sent to South America, Asia and Europe.

PUERTO RICO With no oil or gas production, the country is dependent on petroleum products and natural gas imports to satisfy demand. To decrease fuel costs and reduce emissions, Puerto Rico is shifting from costly fuel oil and diesel to the use of cleaner-burning natural gas for power generation. To accomplish this strategy, Excelerate Energy will build, own and operate the Aguirre Offshore GasPort project, which will be located approximately 4 mi offshore the southern coast of Puerto Rico. The terminal will consist of a floating storage and gasification unit (FSRU), a fixed jetty and a subsea pipeline to deliver imported natural gas to the Puerto Rico Electric Power Authority’s

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Regional Report (PREPA’s) Central Aguirre power plant. With a total capacity of 1,500 MW, the Aguirre power plant is the nation’s largest power facility, and it also has the highest fuel cost of all of PREPA’s facilities. By converting the plant’s feedstock from heavy fuel oil and diesel to natural gas, the nation will save on fuel costs and reduce emissions. The FSRU vessel will have throughput rates of up to 500 MMcfd, with a storage capacity of 3.2

Bcf. The project received US Federal Energy Regulatory Commission (FERC) approval in 2015, and operations are scheduled to commence in 2Q 2017.

COSTA RICA The country’s national oil refiner, Recope, has reinstated plans to expand and modernize the Moin refinery. The project is being developed by the Chinese-Costa Rican Reconstruction Corp. (Soresco), a

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JV between Recope and China National Petroleum Corp. (CNPC). The $1.5-B project, located in the Caribbean port of Limon, will expand the plant’s refining capacity from 25 Mbpd to 65 Mbpd. It will also improve the quality of refined fuels to meet Euro 4 specifications, as well as to produce biofuels. The project was canceled in 2013 by Costa Rica’s comptroller general’s office due to a conflict of interest. The project’s contractors utilized CNPC’s subsidiary, Huanqui Contracting and Engineering Corp., on the feasibility study. This was prohibited in the JV agreement, which stalled the project for nearly two years. However, the project’s feasibility study was revived in 2015 with a new focus on transparency. If greenlighted, the project could create a surge in domestic employment. The expansion and modernization project is expected to create over 2,000 direct jobs and 3,000 indirect jobs during peak construction.

DOMINICAN REPUBLIC In February 2014, Antillean Gas Ltd., a consortium comprising COASTAL (Propagas and Tropigas), Promigas, Ipson, InterEnergy and BW Gas, broke ground on a new LNG receiving terminal. The $350-MM, 1-MMtpy LNG import terminal is located in San Pedro de Macorís, on the country’s southeast coast. The project is part of the country’s plan to import natural gas to fuel power plants and provide power generation for residential, industrial and transportation sectors. The facility, which will process natural gas into more than 1,000 MW of power generation for the country, is expected to be completed in 2016. CUBA The country’s energy sector has been in the doldrums due to limited foreign investments, unsuccessful deepwater drilling, and a lack of oil and gas infrastructure, making it dependent on imports to satisfy fuel demand. The drop in oil prices has also hurt two of Cuba’s biggest investors, Russia and Venezuela. Cuba is struggling to find other partners to invest in its oil and gas sector projects, such as a $6-B expansion and upgrade to the Cienfuegos refinery (FIG. 3). The country has been seeking investors for years to back the plan that would see the Cienfuegos refinery’s capacity expanded

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Regional Report

FIG. 3. The technical renovations needed at Cuba’s four existing refineries, such as the Cupet Cienfuegos refinery located in Cienfuegos, shown here, are stalled for lack of finances.

from 65 Mbpd to 150 Mbpd. The project remains in limbo. Although refining projects have stalled, Cuba claims it will move ahead with nearly $3 B in ammonia-urea and LNG projects. The country is planning to build a $1.4-B LNG regasification plant and a $1.2-B ammonia-urea plant. The LNG import terminal will have a capacity of 2 MMtpy and will provide additional power generation to the island. The ammonia-urea plant will produce 400 Mtpy of urea and 370 Mtpy of ammonia. Excess output is expected to be exported to Central America and the Caribbean. Both of these projects are backed heavily by Venezuela. No timetable for construction or completion has been set. With the drop in oil prices severely affecting Venezuela’s economy, it is unlikely that these projects will be completed soon.

PANAMA The country is not a producer of oil or gas, and it relies primarily on imports to satisfy demand. However, Panama plays a major role in global trade. The Panama Canal is one of the world’s most significant energy transit points, connecting the Pacific Ocean to the Caribbean Sea and the Atlantic Ocean. It allows tanker companies to forego navigating around

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Heat Transfer V. D. SHIRPURKAR and M. E. IBRAHIM, Saudi International Petrochemical Co., Al-Jubail, Saudi Arabia

Calculate thermal efficiency to optimize fired heater operation In a Saudi International Petrochemical Co. petrochemical plant, fired heaters are designed to supply heat duty at 7 MW and are one of the major consumers of energy. Since the esterification reaction to create polybutylene terephthalate (PBT) is endothermic in nature, polymer chain formation is mainly dependent upon the specific heat required to react a large number of molecules to form the PBT product. The tubular heaters use mainly natural gas as fuel gas, which is the main portion of fuel fired in the burners, plus waste organic chemical liquid tetrahydrofuran (LTHF) fuel, which is generated from the tetrahydrofuran (THF) recovery unit. Additionally, offgas is fired into furnaces. However, in case of any upset in process, the offgas is routed to the flare. The burner is designed to fire combined oil and gas fuel, with plant air used to atomize waste liquid fuel. The type of heater used in a PBT plant is forced-draft, which gets its air supply from a combustion air blower. Thermal energy liberated by the combustion of fuel is transferred to stable heat-transfer fluids circulated in tubular coils. The heater is thermally insulated, except for the top side, which is left bare to facilitate access to burner performance checking. This case study examines the operation of the fired heater in a PBT petrochemical unit. The furnace, as shown in FIG. 1, is considered among the best in operation in the country. It is operated with a thermal efficiency of greater than 90%, in compliance with the lowest generation of SOx and NOx. Fired heater efficiency. The successful operation of a fired

heater must aim at the highest possible thermal efficiency together with the lowest pollution. Furnace thermal efficiency is defined as the percent ratio of the total heat absorbed in a furnace to the total heat input supplied: Heat absorbed Heater efficiency = 100 × (1) Heat supplied (NCV+ Qa + Qf ) – Qs – Qr Heater efficiency = 100 × (2) (NCV + Qa + Qf ) Heat absorbed = Total heat supplied – loss from stack – (3) loss from radiation section where: NCV = Net heating calorific value, kJ/kg Qs = Heat loss from stack, kJ/kg

Qr = Heat loss from radiation, kJ/kg Qa = Sensible heat supplied by air, kJ/kg Qf = Sensible heat supplied by fuel, kJ/kg Natural gas supplied to the petrochemical facility is considered as the basis for calculation. Calorific value calculations. The net calorific value (NCV)

and gross calorific value (GCV) of mixed fuel in TABLE 1 is calculated by using published molar calorific values of fuel components. Stoichiometric equations for component combustion are listed in TABLE 5.

Flue gas calculation using combustion reaction. The theoretical or stoichiometric requirement of air required for fuel is calculated by using the combustion reaction with the molar flowrate. The fuel combustion reaction forms water and carbon dioxide and releases heat. The heat capacity of mixtures is calculated in the flue gas analysis shown in TABLE 3.

SAMPLE CALCULATION FOR FUEL REQUIRED To derive the kmol of CH4 required, use Eqs. 4 and 5: 16 kmol of CH4 = 32 kmol of O2

(4)

17.032 kmol/hr of CH4 = (17.032 × 32 ÷ 16) = 34.064 kmol/hr of O2 required

(5)

Similar flowrate calculations for ethane, propane, etc., are shown in TABLE 2. Stack

Process gas; normally no flow Waste THF 35 kg/hr

Furnace

Heat transfer fluid out Heat transfer fluid in

Natural gas 425 Nm3/hr

Combined air 7,157 kg/hr

Flue gas 7,448 kg/hr

FIG. 1. PBT unit furnace flow diagram. Hydrocarbon Processing | MARCH 2016 53

Heat Transfer TABLE 1. NCV calculation Calorific value Components of natural gas feed

Mol%

Nm3/hr

Mol wt

Kg/hr

Kmol/hr

Methane

89.81

Ethane

4.64

381.69

16

272.52

17.032

19.72

30

26.4

Propane

1.72

0.88

7.31

44

14.35

0.326

Butane

0.42

1.785

58

4.62

0.08

Nitrogen

2.56

10.88

28

13.59

0.485

Isobutane

0.3

1.275

58

3.3

0.057

CO2

0.45

1.9125

44

3.76

0.085

Hydrogen sulfide

0.1

0.425

34

0.64

0.019

Total natural gas supply

100

425

–

339.18

18.96

Component waste, liquid tetrahydrofuran (THF)

Mol wt

Kg/hr

Kmol/hr

THF

72

10

0.139

Butanol

74

25

0.338

374.18

19.44

Total fuel flow (LTHF + natural gas), kg/hr

Kmol SO2 formed

Kmol CO2 formed

Kmol H2O formed

34.064

0

17.032

34.064

Ethane

3.08

0

1.76

2.64

Propane

1.63

0

0.326

0.652

Butane

0.52

0

0.32

0.4

30

Isobutane

0.37

0

0.228

0.285

20

Hydrogen sulfide

0.028

0.019

0

0.019

10

Tetrahydrofuran

0.764

0

0.55

0.556

Butanol Total, kmol/hr

45,630 kJ/kg

70 Excess air, %

60 50 40

0 0

2.02

0

1.35

1.689

42.482

0.019

21.571

40.303

TABLE 3. Components of stack flue gas flow Component

50,417 kJ/kg

80

Kmol O2 required

Methane

Net

90

TABLE 2. Flow rates for oxygen, reactants and products Component

Gross

Kmol/hr

Mol wt

Kg/hr

O2

9.346

32

299.1

N2

195.5

28

5,473

CO2

21.57

44

949.1

SO2

0.019

64

1.216

H2O

40.3

18

Total flue flow at stack, kg/hr

Heat capacity, kJ/kg°K

1.1 (simulation calculation)

725.5 7,448

Stoichiometric, or theoretical, requirement of air. The stoichiometric, or theoretical, requirement of oxygen is 42.482 kmol/hr. Oxygen content on a flue gas dry basis = 4%, based on an oxygen analyzer reading. Refer to FIG. 2 for the percentage of excess air supply as a general rule in industrial practice. Eqs. 6 and 7 can be used to calculate the stoichiometric, or theoretical, requirement of air:

Actual O2 supply in kmol/hr = (1 + 0.22) × 42.482 = 51.828 54 MARCH 2016 | HydrocarbonProcessing.com

(6)

5

O2 content in flue gas

10

15

FIG. 2. Flue gas oxygen content vs. excess air.

Actual air supply in kmol/hr for 51.78 kmol of O2 = (100 × 51.828) ÷ 21 = 246.802 kmol/hr

(7)

Calculating nitrogen in flue gas. Since nitrogen does not take part in the combustion, but is the major component in air—in addition to oxygen, minor inert gases, CO2 and moisture—it is a significant component in emitted flue gas. Eq. 8 shows how the nitrogen flowrate is calculated: O2 in flue gas = actual O2 supplied – actual O2 (8) used in combustion Actual O2 supplied = 51.828 kmol/hr Actual O2 used in combustion = 42.482 kmol/hr O2 in flue gas = 9.346 kmol/hr Nitrogen in flue gas = (actual air supplied – actual O2 supplied) + N2 in fuel = (246.802 – 51.828) + 0.48 = 195.459 kmol/hr Calculating sensible heat of air Qa. How much heat is

contributed by combustion air depends on the level of inlet air preheat. Eq. 9 shows how sensible heat of air is calculated: Qa = M × Cp × (air inlet temp – datum temp)

(9)

Heat Transfer where: M is the flow rate and Cp is the specific heat = 7,157.245 × 1 × (30 – 15) = 107,359 kJ/hr = 107,359 ÷ 374.18 = 286.9 kJ/kg

TABLE 4. Furnace emissions against the national norms Emissions summary

Calculating sensible heat of fuel Qf. See Eq. 10 for cal-

culating the contribution to sensible heat and precombustion, based on inlet fuel temperature: Qf = M × Cp × (fuel inlet temp – datum temp)

(10)

= 374.18 × 2.07 × (30 – 15) = 11,618.289 kJ/hr = 11,618.289 ÷ 374.18 = 31.05 kJ/kg

NOx, tpy

SO2, tpy

Standard

9.37

Actual value

7.89

230

0.03

180

CH4 + 2 O2 = CO2 + 2 H2O 2 C2H6 + 7 O2 = 4 CO2 + 6 H2O

2 C4H8O + 11 O2 = 8 CO2 + 8 H2O C4H10O + 6 O2 = 4 CO2 + 5 H2O

(11)

Flue gas density at 1 bar pressure

1.4 1.2

(12)

Density, kg/m3

1.0 0.8 0.6 0.4 0.2 0.0

0

50

100

150 Temperature, °C

200

300

400

FIG. 3. Flue gas density vs. flue gas temperature.

Calculation of heater efficiency. See Eq. 13 for this calcu-

lation using the same sample data:

(13)

(45,630 + 286.91 + 31.05–3,612.6 – 912.6 = 100 45,630 + 286.91 + 31.05 = 90.15%

EMISSIONS SUMMARY CALCULATION AND REGULATORY NORMS The basis of calculation is considered as furnace operation based on 8,000 hr over the entire year. The online analyzer shows an average reading of CO at 0.3 ppm and NOx at 48.40 ppm. Density of flue gas is shown in FIG. 3. At 180°C, flue gas density equals 0.75 kg/m3, calculated using simulation software. The flue gas volumetric flowrate in m3/hr equals 7,447.36 ÷ 0.75 = 9,929. CO produced. See the Eq. 14 calculation of flowrate for the

sample data set: mg/m3 = PPM × mol wt ÷ 22.41 = 0.30 × 28 ÷ 22.41 = 0.375 mg/hr = mg/m3 × m3/hr (volumetric flowrate)

7.81

9

H2S + 1.5 O2 = SO2 + H2O

= 45,630 × 0.02 = 912.6 kJ/kg

(NCV + Qa + Qf ) – Qs – Qr NCV + Qa + Qf

21.2

C3H8 + 5 O2 = 3 CO2 + 2 H2O

Heat loss by radiation Qr. The radiation heat loss is generally considered to be between 2% and 4%. This is dependent upon material used in the insulating refractory lining and the thickness of insulation. The refractory is composed of alumina and silicate-based low-expansion cement cast to withstand a maximum recommended temperature of up to 1,650°C. Note that Eq. 12 uses 2% for radiation heat loss.

Heater efficiency = 100

Stack outlet temp, °C

TABLE 5. Combustion component reaction equations

= 7,447.735 × 1.1 × (180 – 15) = 351,764 kJ/hr = 351,764 ÷ 374.18 = 3,612.6 kJ/kg

Qr = Net calorific value × 0.02

CO, tpy

C4H10 + 6.5 O2 = 4 CO2 + 5 H2O

Calculating heat loss by stack Qs. See Eq. 11 for calculating flue gas heat loss up the heater stack:

Qs = M × Cp × (flue gas outlet temp – datum temp)

Pollutants

(14)

= 0.375 × 9,929 = 3,723 kg/hr = 0.0038, tpy = 0.03 NOx produced. See the Eq. 15 calculation of flowrate:

mg/m3 = ppm × mol wt ÷ 22.41 (15) = 48.40 × 46 ÷ 22.41 = 99.3 mg/hr = mg/m3 × m3/hr (volumetric flowrate) = 99.3 × 9,929 = 985,950 kg/hr = 0.985, tpy = 7.89 TABLE 4 compares emissions against the national norms; this fired heater performs significantly better than the government’s regulatory requirements. VIJAY D. SHIRPURKAR is a senior process engineer at Saudi International Petrochemical Co. in Jubail, Saudi Arabia. He has 13 years of technical and plant operational experience in the refining and petrochemical sectors. He has worked with refinery delayed cokers in the world’s largest grassroots refinery (Reliance industries Ltd.,’s Jamnagar plant in India) as a manager for LPG and fuel gas separation units. He is a graduate chemical engineer from the Dr. Babasaheb Ambedkar Technological University in Lonere, India. MOHAMMAD E. IBRAHIM is the general manager of operations at Saudi International Petrochemical Co. in Jubail, Saudi Arabia. He started his career in process engineering and gained wide operational and troubleshooting experience in working for international gas companies, especially in the area of reformer furnaces. He graduated in chemical engineering from King Fahd University of Petroleum and Minerals in Dhahran, Saudi Arabia. Hydrocarbon Processing | MARCH 2016 55

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Heat Transfer C. BAUKAL and B. JOHNSON, John Zink Hamworthy Combustion, Tulsa, Oklahoma; and R. NEWNHAM, OnQuest Canada ULC, Calgary, Alberta, Canada

Minimize unplanned shutdowns of fired heater operations There are many potential “rules,” or guidelines, for the safe operation of process heaters. American Petroleum Institute (API) STD 5601 and API RP 5352 provide many useful recommendations for operating heaters and burners, respectively. Companies normally develop their own detailed procedures based on API recommendations and their own experiences. These best practices can be summed up in four simple and easyto-remember rules for operating fired process heaters. These rules are intentionally broad to encompass many of the more detailed procedures that have been developed for a particular heater, and they are especially useful for new operators that are learning the many facets involved with safely running process heaters. These common-sense guidelines are not intended to replace company procedures, but to provide a framework to more easily remember some of the more important factors in the safe operation of fired heaters. The focus here is upon safety,3 and not specifically, for example, on minimizing pollution,4, 5 or in maximizing efficiency6, 7 or uptime. However, these are often corresponding results of following the principles presented here. Examples are also provided for the potential negative consequences of not following each rule.

RULE 1: KEEP THE FLAMES IN THE BOX Keeping the flames inside the firebox seems to be an obvious rule since the goal is to heat something inside the heater, not outside. This is vital for both worker safety and equipment integrity, and includes keeping not only flames in the heater, but also the hot gases generated in the combustion process. Flames and hot gases could exit any heater openings, including sight ports, burner air inlets and cracks in the heater shell. FIG. 1 shows an example of a flame outside a heater. In this case, the flame periodically exited the heater when the pressure inside went positive, which was on a fairly regular basis. Hot gases exiting a heater may or may not be visible, depending on the conditions. On a bright sunny day, hot gases that are exiting heater openings might be difficult to see, so it is critical to ensure that no conditions exist that could force flames and hot gases outside a heater. The high-temperature gases can injure personnel, even those wearing flame-retardant clothing. Prolonged emissions of hot gases can also damage equipment. Two common causes that can force flames and hot gases out of a heater are positive pressure and pulsations, which are both discussed next.

Heater draft. Process heaters are designed to have a slightly

negative pressure at the top of the radiant section, which is also referred to as the arch, or bridgewall.8 FIG. 2 shows a typical process heater draft profile. The least-negative (lowest-draft) location inside the heater is at the arch. However, it is possible for the pressure there to go positive. The primary device used to control heater draft is the stack damper, which, if not sufficiently open, would cause the pressure to go positive. A plugged convection section could also cause the same effect. Note: It is possible that the heater pressure could be negative near the bottom of the heater, yet positive near the arch. This is why the heater pressure should be checked before attempting to look through the sight ports, especially those that are located where the pressure is most likely to be positive if the heater is not operating as designed. Draft measurement at the arch is essential, and automatic draft control is recommended to ensure the heater pressure does not go positive. Flame stability. While there are many types of unstable flames, those that are pulsing—often referred to as huffing or woofing—can force flames and hot gases to go outside of a heater.9 The transient nature of this condition can cause the

FIG. 1. Flames and hot gases can escape heater openings, including sight ports, burner air inlets and cracks in the heater shell. Hydrocarbon Processing | MARCH 2016 57

Heat Transfer pressure in a heater to fluctuate between positive and negative. Severe pulsing has been known to cause sight port covers and explosion doors to lift or flap. Pulsating flames can cause flames and hot gases to escape a heater, and they can also potentially cause flames to temporarily escape by lifting them completely off the burner tip. If there is an ignition source inside the heater, or in any part of the heater that is at a temperature above auto-ignition, flames could reignite.10 This is potentially very dangerous, as the heater would likely be full of flammable gases that could ignite explosively, causing severe over-pressurization. One of the possible causes of pulsing flames is the fuel pressure to the burners exceeding the maximum design pressure.8 High fuel pressure increases the fuel/air mixture velocity exiting the burner. If that velocity significantly exceeds the burning velocity, the flames can lift and even go out completely. The flames become over-strained because the gases cannot react fast enough to maintain flame stability. In diffusion burners, also known as raw gas or nozzle-mixed, the mixing of the fuel and air begins at the exit of the burner. Higherthan-designed fuel exit velocities can also delay the mixing process, which can increase the likelihood of flame instabil120 End of stack 100

After damper

60

After convection

raft

20

gn d

Floor -0.5

RULE 2: KEEP THE FLAMES OFF THE TUBES Fired heaters are used in the refining and petrochemical industries for the heating of various fluids. FIG. 3 shows heat that is transferred by both radiation and convection to the outside of the process tubes, through tubes by conduction, and away from the inside of the tubes by convection. Flame impingement causes too much heat transfer to the outside of the tubes.

40

Arch

-0.6

Elevation, ft

80

ity. Before the flame blows off completely, it often begins to pulse. This is because the high-velocity gases slow down as they expand after exiting the burner, allowing them to burn back toward the burner. Assuming that the gas velocities are not so high that the flame blows off completely, this can set up a pulsation where the flame burns back toward the burner but then is pushed away by the higher-than-designed gas velocities. The pulsating heater pressure can cause flames and hot gases to exit the heater. This cycle continuously repeats until something changes. There are other possible causes of flame pulsations, such as damaged flame stabilizers (e.g., metal-cone flame holders or ceramic tile ledges) or a problem with the fuel injection system.9 Flame stabilizers are designed to anchor the flames close to the burner outlet. Damaged stabilizers can lead to unstable flames. There are several possible fuel injection system problems that could produce unstable flames, such as tips that are plugged or damaged, improperly aligned or the incorrect size. The proper installation and maintenance of the correct burner tips are essential for proper burner operation and pulsation prevention.

Desi

-0.4 -0.3 -0.2 -0.1 Pressure, in. of water column

0

0.1

FIG. 2. A typical process heater draft profile, which is designed to have a slightly negative pressure at the top of the radiant section, or arch.

Conduction–heat transfer through steel tube Convection section

Radiant section

Radiation–heat from flame and hot walls Convection-flowing hot furnace gases

FIG. 3. Heat is transferred by both radiation and convection to the outside of the process tubes.

58 MARCH 2016 | HydrocarbonProcessing.com

FIG. 4. As most process heaters have a hydrocarbon feed, flame impingement can cause serious problems.

Heat Transfer Flame impingement on process tubes. Since most process heaters have a hydrocarbon feed, flame impingement can cause serious problems. Visual observation of the burners may show that the flames are contacting the tubes. In some cases, a gradual increase in the tube metal temperature (TMT) could also indicate possible flame impingement. Operators should make it a point to look into each of their fired heaters at least once a shift to check for any problems with the flame patterns, particularly flame impingement (FIG. 4). The effect on operations. The reason that tubes do not overheat inside a furnace is because of the cooling effect of the process fluid inside the tubes. This is the reason many heaters have carbon steel tubes: once a tube begins to overheat, there is a gradual buildup of carbon on the inside of the tube. This insulates a tube from the cooling effects of the process flow, which, in turn, makes the tube hotter. As the carbon continues to build, the flow area of the tube is reduced. If allowed to continue, the carbon will choke the tube completely, potentially resulting in a tube rupture. Hot spots will normally develop in progressive stages. When the flames contact the tube surface, there is a cooling effect on the flame. This results in ash being laid down on the tube. This buildup will lead to scale on the tubes as the outer layer of the tube starts to burn away. There are various stages in this process, illustrated in FIG. 5:11 • Dark areas first begin to appear from the carbon coating on the side of the tubes facing the burners. • Silver or light gray spots form within the dark areas. This is caused by the carbon being burned off. • These light gray spots will enlarge and cover a larger area. • As the coking continues, red spots will begin to appear in the gray areas of the tubes. In some cases, the tube will take on a “mirror” finish that looks almost like a chromed piece of pipe. • The tube will eventually start to bulge and then develop “pinhole” leaks. The tube is ready to rupture, and immediate action must be taken. Preventive and corrective actions. Keeping the flames off

the tubes is paramount. If flame impingement is noticed, the first step should be to adjust the burner causing the impingement to get the flame off the tube(s). • The burner air register should be checked to confirm it is open. Gas tips should be examined for any plugging that could cause the flame to impinge on the tube. If plugging is evident, the tips must be removed and cleaned. Gas tips can be properly orientated by checking the burner drawing. • Oxygen and draft requirements as per the heater design specifications should be confirmed. If the heater cannot be shut down, there are three options: 1. Take the burner out of service, or reduce the firing rate by manually closing the block valve. 2. Increase the excess air to help cool the firebox. 3. Increase the process flow to the overheated pass. There are other options, such as wrapping the tubes or clamping them while the heater is in service, which must be considered as extreme risk situations and approved by facility safety experts.

RULE 3: KEEP THE PROCESS IN THE TUBES Generally, the purpose of a process heater is to heat some type of fluid flowing through tubes inside the heater. In a refinery or chemical plant, the fluid is typically some type of hydrocarbon fluid, such as crude oil. The burners in the heater then heat the fluid. More accurately, the burners heat the tubes, which conduct the heat into the fluid flowing through them. The tubes are used to safely convey the fluid through the heaters. If those tubes are damaged, the fluid that is under pressure can leak out into the heater. This is particularly dangerous when the fluid is flammable because the heater is likely to be hot enough that flammable fluid could ignite if there is sufficient oxygen in the heater. There are many possible causes for a tube leak. The tubing could have been improperly manufactured with some thinner sections that fail under pressure or prematurely. It has become more common to inspect brand new tubes before they are installed in a heater with some type of tube inspection system to ensure that this does not happen. Tube leaks can also be caused by the presence of a corrosive process fluid (FIG. 6). An alternate metal type may be required if the tubing is likely to fail due to corrosion before it can be replaced at the next turnaround. An-

FIG. 5. Hot spots in tubes develop in progressive stages, leading to scaling as the outer layer begins to burn away.

FIG. 6. Corrosive fluids can cause a pipe to rupture. Hydrocarbon Processing | MARCH 2016 59

Heat Transfer other potential cause for a tube leak is operating a heater with insufficient process fluid flow, which is required to remove heat from the heater to keep the process tube temperature below its design limit. This may happen at startup if there are problems getting the flow established. It could also be caused by a tube failure or process flow valve being closed upstream of the heater. As related in Rule 2, the most common cause for damaged process tubes is flame impingement. Burners that are located too close together can cause the flames to coalesce and create a much larger and longer flame that could impinge on tubes near or above the burners (FIG. 7). Continuous flame impingement on a process tube causes the tube wall temperature to increase. If the temperature gets high enough, it can cause the hydrocarbon fluid flowing through the tube to break down and deposit

coke (carbon) on the inside of the tubes. Tubes are designed to be cooled by the fluid flowing through them by transferring heat away from the metal by internal forced convection. However, if coke forms inside the tube, it acts like an insulator since its thermal conductivity is much lower than the metal’s. This reduces the convective cooling of the metal, causing the rise in the metal temperature. When internal coking is coupled with flame impingement, the heater tube temperature can continue to rise above its maximum design limit and eventually cause the tube to leak or even rupture. This must obviously be avoided— a significant tube leak can generate a large fire and huge quantities of thick black smoke as there will most likely not be enough oxygen to fully combust all of the leaking flammable fluid. A further concern is that the uncombusted leaking liquids and gases could burn outside the heater, increasing the risk of personnel injury and equipment damage.

Convection section process tubes

RULE 4: KEEP FLAMMABLES OUT DURING LIGHTOFF Although this is listed as the fourth rule, purging is the first action in the safe startup of all direct-fired heaters.9 Therefore, purging is a mandatory requirement and is covered in all the codes and standards around the world, such as API 556; NFPA 85, 86 and 87; EN 298; CSA B149.3–10; and ANZ 3814. Depending on the jurisdiction and authority under which the heaters are operating, the purging function is not only mandatory, but it must also be a “proven” function within the safety system operating the heater.

Flame starved for air

Flame-flame interaction Radiant section process tubes Recirculation zone Tight burner circle

FIG. 7. Continuous flame impingement on a process tube causes the tube wall temperature to increase.

Stack exit

Natural purging. In the past, and in many installations today, the normal practice of purging a natural-draft heater is to alStack exit

Stack Stack damper Stack transition Convection section

Radiant section

Burners FIG. 8. Ambient air passes through a heater by natural draft, purging the heater of any fuel products.

60MARCH 2016 | HydrocarbonProcessing.com

Stack damper Stack transition Convection section

Radiant section

Steam purge/snuffing connections Burners FIG. 9. Injecting steam into the base of the radiant section pushes any combustible gases up while bringing in fresh air through the burner air registers or dampers.

Heat Transfer low time for ambient air to pass through a heater by natural draft, thus purging the heater of any fuel products (FIG. 8). Typically, this takes about 20 minutes on a cold start, and is left to the discretion of the operator. The stack damper and burner air registers must be in the open position. The major drawback with this purging method is proving that the purging function has taken place and is complete. The ambient conditions have a big effect on the freeflow of air through a heater. In colder regions, the freeflow of air can be zero, so there is no actual purging taking place. For that reason, ambient purging based purely upon time does not guarantee that a heater has actually been purged, and, therefore, should not be done. The time designation of 20 minutes is not based on any empirical calculated method to ensure that four or five volume changes take place, but this is an industry predetermined minimum time period. Most operating companies will supplement the 20-minute time period with an operator-actioned lower explosive limit (LEL) check for any combustibles in the floor of the firebox. Steam purging. An alternative to using ambient air is to purge with steam (FIG. 9), which involves injecting steam into the base of the radiant section. As this hot steam rises, it pushes any combustible gases up through the heater while bringing in fresh air through the burner air registers or dampers. However, caution must be taken to prevent steam condensation

Stack exit Stack Steam injection Stack damper Stack transition Convection section

Radiant section

Burners FIG. 10. Injecting steam into the stack above the damper and utilizing a steam educator draws air into the bottom of the heater and pushes it up, expelling any combustible gases.

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Hydrocarbon Processing | MARCH 2016 61

Heat Transfer Manual checks complete

Purge permissives satisfied

And

Ready to purge

Start purge

Start purge fan

Start 5-minute purge timer

Purge flow detected in 1 minute?

Yes

Alarm reset

No

Common alarm Purge fail

Purge in progress Purge permissive OK and flow maintained?

No

Timer complete? Yes

LITERATURE CITED ANSI/API Standard 560, “Fired heaters for general refinery service,” Fourth ed., 2007, American Petroleum Institute, Washington, DC. 2 API Recommended Practice 535, “Burners for fired heaters in general refinery services,” Third ed., 2014, American Petroleum Institute, Washington, DC. 3 Baukal, C. E., “Safety,” The John Zink Hamworthy Combustion Handbook, Vol. 2: Design and Operations, CRC Press, Boca Raton, Florida, 2013. 4 Baukal, C. E., I-P Chung, S. Londerville, J. G. Seebold and R. T. Waibel, “Pollutant Emissions,” The John Zink Hamworthy Combustion Handbook, Vol. 1: Fundamentals, CRC Press, Boca Raton, Florida, 2013. 5 Baukal, C. E. and W. Bussman, “NOx emissions,” The John Zink Hamworthy Combustion Handbook, Vol. 1: Fundamentals, CRC Press, Boca Raton, Florida, 2013. 6 Baukal, C. E. and W. Bussman, “Thermal Efficiency,” The John Zink Hamworthy Combustion Handbook, Vol. 1: Fundamentals, CRC Press, Boca Raton, Florida, 2013. 7 Newnham, R., Direct-Fired Heaters: Improving Efficiency and Capacity While Reducing Emissions, Kingsley Knowledge Publishing, Alberta, Canada, 2013. 8 Waibel, R. T., M. G. Claxton and B. Reese, “Burner design,” The John Zink Hamworthy Combustion Handbook, Vol. 2: Design and Operations, CRC Press, Boca Raton, Florida, 2013. 9 Newnham, R., Direct-Fired Heaters: Operator Training Manual, Kingsley Knowledge Publishing, Alberta, Canada, 2013. 10 Newnham, R., Direct-Fired Heaters: A Practical Guide to Their Design and Operation, Kingsley Knowledge Publishing, Alberta, Canada, 2012. 11 Johnson, W., E. Platvoet, M. Pappe, M. Claxton and R. Waibel, “Burner Troubleshooting,” The John Zink Hamworthy Combustion Handbook, Vol. 2: Design and Operations, CRC Press, Boca Raton, Florida, 2013. 1

Purge in progress

Yes No

These simple and commonsense rules for the safe operation of a process heater—keeping the flames in the box and off the tubes, maintaining the process in the tubes, and shutting flammables out during lightoff—are not intended to be comprehensive, but are designed to be easy to remember, particularly for those with less experience in operating process heaters. Failure to abide by these four rules can produce serious consequences, including personnel injury and equipment damage. Companies should have their own detailed processes and procedures for operating specific heaters. Heaters must be constantly monitored with regular visual inspections both inside and outside of the heater. This will help prevent flames from exiting the heater, along with the leakage of flammable liquids from the process tubes.

Stop purge fan Purge complete

To light pilots FIG. 11. A typical flow diagram for a mandatory purge using an automated burner management system.

from affecting the operation of electronics in or near the heater, such as the ignition and flame monitoring instrumentation. Another way to purge using steam is with a steam educator (FIG. 10). Steam is injected into the stack above the damper. As the steam rises, it draws air into the bottom of the heater and pushes it up through the heater, expelling any combustible gases. Mandatory purging with a burner management system. The most important function for the use of a burner

management system (BMS) is to prevent the possibility of an accumulation of combustible gas within the heater prior to introducing a flame inside the heater.10 The issue becomes how to “prove” that four to five volume changes have taken place. With a steam purge, or a purge using a steam educator, this is possible. If it is proven with the stack damper open, and if the flowrate of the purge steam is known, it can be determined how long it takes to evacuate the air and any combustibles from within the heater. With a natural purge, this is not possible. For facilities that lack available steam, the next resource is to install a purge fan to perform the purge function. With this method and the known flowrate of air, the purge time period can be determined. FIG. 11 shows a typical flow diagram for the mandatory purge using an automated burner management system. 62 MARCH 2016 | HydrocarbonProcessing.com

CHARLES BAUKAL is the director of the John Zink Institute, which is part of John Zink Co. LLC, where he has worked since 1998. He is the author and editor of 13 books on industrial combustion, including multiple versions of The John Zink Hamworthy Combustion Handbook, and is also an inventor on 11 US patents. Dr. Baukal has 35 years of industrial experience, and earned BS and MS degrees in mechanical engineering from Drexel University, an Ed.D from Oklahoma State University, and a PhD in mechanical engineering from the University of Pennsylvania. He also holds numerous professional engineering and environmental engineering licenses. BILL JOHNSON began his career in the power industry in 1969 designing coal-fired power plants. He started working at John Zink in 1977, and his roles have included the design, sales, testing and startup of process burners. Mr. Johnson is a co-author of The John Zink Hamworthy Combustion Handbook. He has written several technical papers for the hydrocarbon processing industry (HPI), and has been involved with the John Zink Burner School since 1982. ROGER NEWNHAM is the senior vice president for OnQuest Canada ULC, a Primoris company. He has over 40 years of experience in the design and operation of direct-fired heaters, including the writing and publishing of four books on this subject. He earned a BSc degree in mechanical engineering from Brighton University in the UK, and is a fellow of the Institute of Mechanical Engineers.

Maintenance and Reliability K. R. RAMAKUMAR, Johnson Matthey Process Technologies, UAE

Detect boiler leaks upstream of the shift reactor in H2 plants In a typical hydrogen (H2 ) flowsheet, water gas shift (WGS) reactors are normally located after the steam reforming section. Common products of steam reforming include carbon monoxide (CO), which the WGS reaction converts to produce additional H2 , as shown in Eqs. 1 and 2, making WGS reactors important for increased H2 production: CO + H2O = CO2 + H2 ; ΔH = –41.1 kJ/mol

(1)

The reaction is exothermic, and high conversions are favored by low temperature and high steam-to-dry-gas ratio. Ammonia plants usually operate a two-stage system; a hightemperature shift (HTS) followed by a low-temperature shift (LTS), with a suitable form of inter-bed cooling. H2 plant designs feature a number of differing shift conversion sections. There is commonly an HTS stage followed by a pressure-swing absorption (PSA) unit to separate H2 product from other components. Occasionally, a medium-temperature shift (MTS) is used in preference to an HTS, with the main driver for an MTS being a lower steam-to-carbon (S:C) ratio, which is associated with higher process efficiency. In flowsheets where an HTS is in place, the S:C ratio cannot be less than 2.8. In older H2 plants, a two-stage system is often utilized in which an HTS is followed by an LTS stage (with suitable interbed cooling), CO2 removal and, finally, the methanation stage. There are also a few plants with an HTS-LTS-PSA configuration for the purpose of producing more H2 and using a smaller PSA unit. The evolution of different configurations is shown in FIGS. 1, 2 and 3. The HTS reactor typically operates at an inlet temperature range of 320°C–360°C; an “exotherm” equivalent to a temperature rise of 55°C–65°C is normally encountered. MTS catalysts tend to run at an inlet temperature range of 200°C–230°C, with

an exotherm equivalent to a temperature rise of 70°C–100°C. LTS and MTS catalysts operate close to condensation conditions during the early part of the catalyst life. LTS catalysts typically operate at an inlet temperature range of 190°C–220°C, with an exotherm equivalent to a temperature rise of 15°C– 30°C. Typical shift reactor parameters are shown in FIGS. 4 and 5. The inlet temperature of an LTS/MTS should be at least 15°C above the dew point at all times to avoid capillary condensation in the pore structure. Condensation and liquid water in the shift reactor can adversely affect the life of the catalyst by washing soluble catalyst poisons through the bed. In addition, there is a possibility that wet catalyst may be damaged by the rapid evaporation or boiling of water within the pellets if it is heated too quickly during a plant restart. Although current HTS/MTS/LTS catalysts are quite robust and can withstand PSA offgas

From steam reforming

HTS

H2O

FIG. 2. A configuration of an H2 plant (since the mid-1980s). PSA offgas

From steam reforming

HTS

HTS

CO2 removal

LTS

Hydrogen product

PSA

H2O

CO2 to vent

From steam reforming

Methanation

Hydrogen product

From steam reforming

PSA offgas

MTS

PSA

Hydrogen product

H2O

H2O

FIG. 1. A configuration of an H2 plant (typical of the 1970s).

Hydrogen product

PSA

FIG. 3. Current H2 plant design options. Hydrocarbon Processing | MARCH 2016 63

Maintenance and Reliability wetting, to some extent, prolonged wetting can physically damage the catalyst, causing a higher pressure drop. Another issue that cannot be neglected is any boiler leak upstream of an HTS/MTS. The process gas boilers are mostly fire-tube boilers that are designed to cool the reformed process gas exiting the reformer. In such boilers, process gas enters the tubes and heats the boiler feed water (BFW) as it enters the shell through downcomer pipes from the steam drum. A twophase steam water mixture within the riser lifts itself back into the steam drum due to the density difference between the incoming water and outgoing two-phase mixture. Based upon the location of a BFW leak, there could be different scenarios by which the downstream shift catalyst could be affected: 310°C–370°C Process gas from SMR 800°C–930°C

190°C–220°C 10–15 mol% CO

2–4 mol% CO

HTS ∆T = 55°C–65°C

LTS ∆T = 15°C–30°C

Gas to PSA/ methanator

Steam generation Steam generation/feed preheat

0.1–0.3 mol% CO

FIG. 4. Typical parameters in an HTS-LTS configuration.

Leak towards outlet side (steam side) • No condensation case: There is a higher ΔP across the shift catalyst bed due to an increased mass flow. • Condensation case: Apart from a higher ΔP due to the above reason, there is a possibility of physical damage of the catalyst due to re-vaporization of condensate, leading to a further increase in ΔP and pushing the poisons farther down the bed. Leak towards inlet side (water side) • Complete vaporization before reaching the catalyst bed: In addition to the increased ΔP contributed by increased mass flow, the residual solids present in the water can contribute to farther increase in ΔP. FIG. 6 shows a classic example of such a case. • Incomplete vaporization before reaching the catalyst bed: In addition to the effect above, the catalyst can also be subjected to physical damage due to the re-vaporization of condensate, leading to a further increase of ΔP and pushing the poisons further down the bed. In any case, the boiler leak could be detected by performing an overall mass balance and observing the temperature profile across the shift catalyst. As shown in FIG. 6, the pressure drop slowly builds up at the onset of boiler feed water ingress into the HTS reactor. It continues to build up slowly for about a year, after which the catalyst pellets at the bottom start to crush as the hydraulic load surpasses the strength of the pellets. TABLE 1. Steam reformer inlet and outlet conditions

2016 WOMEN’S

Reformer inlet conditions Molar flow of feed gas

WGLConference.com

Mass flow of feed gas S:C ratio

131.7 TPH

Total mass flow at reformer inlet

176.2 TPH

Inlet pressure

33 barg

Inlet temperature

540°C

H2

November 1–2, 2016

Hyatt Regency Houston / Houston, Texas 64 MARCH 2016 | HydrocarbonProcessing.com

44.5 TPH 3.0 mol/mol

Mass flow of steam

Typical reformer feed composition

Save the Date

2,500 kmol/h

mol% wet 1.27

CO2

1.28

H 2O

74.51

C1

21.66

C2

0.90

C3

0.25

C4

0.04

C5

0.04

C 6+

0.05

Reformer outlet conditions Total molar flow

13,965 TPH

Total mass flow

176.2 TPH

Pressure

31.3 barg

Outlet temperature

900°C

Maintenance and Reliability Therefore, it is very important for the plant process and operation engineers to detect boiler leaks well in advance to avoid the deterioration of shift catalyst integrity. One of the key practices that helps identify the upstream boiler leak is to perform a mass balance across the shift reactor. Additionally, it is a good practice to keep track of the condensate flow to the deaerator or stripper. Examples of sample mass balances that help detect boiler leaks upstream of the shift reactor are illustrated in TABLE 1, which shows reformer and HTS compositions; TABLE 2, a normal case with no BFW leak; and in TABLE 3, which has a leak. In the mass balance example, it has been assumed that the BFW leak increases the inlet steam:dry gas ratio from the

normal value of 0.493 to 0.60. Based upon this assumption, it can be seen that almost 18% (mass flow) excess condensate is produced. Where there is a provision for separate sample points, it is advisable to take two samples: one immediately downstream of the reformer, and the other at the inlet of the HTS/MTS reactor, downstream of the waste heat boiler (WHB). By performing a wet compositional analysis of these two samples, it can be well established if there is a BFW leak. In plants where there is no provision for collecting samples at two locations, it is suggested instead to observe the daily trend of condensate flow from catch pot to deaerator/strip60

200°C–230°C Process gas from SMR 800°C–930°C

Start of bed crushing

50 Pressure drop, psi

10–15 mol% CO

MTS ∆T = 70°C–100°C

Start of persistent boiler leak

40 30 20

Hydraulic load on bottom catalyst > catalyst strength; therefore, catalyst at bottom collapses

10

Gas to asborber/ methanator

0

0

5

10

15

0.8–1.0 mol% CO

Steam generation

FIG. 5. Typical parameters in an MTS.

20 25 Time online, months

30

35

40

45

FIG. 6. The impact of boiler leak on an HTS catalyst with residual solid buildup.

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Hydrocarbon Processing | MARCH 2016 65

Maintenance and Reliability 426

TABLE 2. HTS reactor inlet and outlet conditions NORMAL CASE (no leaks) mol% wet

mol% dry

406 396

H2

48.63

72.62

CO

10.44

15.59

CO2

5.53

7.96

H2O

33.02

–

2.56

3.83

C1 Steam:dry gas ratio

Temperature, °C

Typical composition at reformer outlet/inlet to HTS reactor

416

386 376 366

0.493

Typical composition at HTS reactor outlet

356

mol% wet

mol% dry

H2

56.13

75.38

CO

2.94

3.95

CO2

12.83

17.22

H2O

25.53

–

C1

2.56

3.44

Steam:dry gas ratio

0.343

HTS outlet molar flow

13,962 kmol/h

HTS outlet mass flow

176.2 TPH

Assuming 99% condensate recovery at catch pots, recovered condensate flow = 0.99 × 13,962 × 0.2553

3,529 kmol/h 64 TPH

TABLE 3. Boiler feed water leak mass balance case Boiler feedwater leak case Inlet to HTS reactor

mol% wet

mol% dry

45.39

72.62

CO

9.74

15.59

CO2

4.98

7.96

H2

H2O

37.5

–

C1

2.39

3.83

Steam:dry gas ratio

0.60

Typical composition at HTS reactor outlet

mol% wet

mol% dry

H2

52.86

75.38

CO

2.26

3.95

CO2

12.45

17.22

H2O

30.03

–

C1

2.39

Steam:dry gas ratio

3.44 0.429

HTS outlet molar flow

14,258 kmol/h

HTS outlet mass flow

185 TPH

Assuming 99% condensate recovery at catch pots, recovered condensate flow = 0.99 × 14,258 × 0.3003

66 MARCH 2016 | HydrocarbonProcessing.com

4,239 kmol/h 76 TPH

S:G 0.6 S:G = 0.493 S:G 0.75

346 336 0

1

2

Bed depth, m

3

4

5

FIG. 7. An HTS catalyst bed temperature profile for different S:G ratios

per. Another way of establishing a boiler leak when there is no separate sample point at the inlet of an HTS/MTS is by using common sense and observing the compositions of the reformer outlet and HTS/MTS outlet. The water gas shift reaction is an equimolar reaction, the excess moles of H2 produced at the outlet of the HTS/MTS should ideally be equal to the moles of water (H2 O) consumed. If the moles of the produced H2 (outlet/inlet) are invariably greater than the moles of the consumed H2 O (inlet/outlet), then there is a possible BFW leak upstream of the HTS/MTS that is responsible for the additional H2 O moles at the outlet of the HTS/MTS. Since CO conversion is favored at high steam:dry gas ratios, the CO at the outlet of an HTS/MTS would also be less than expected. The bed temperature trend would also help in complementing the above observation. With a leak, the exotherm across an HTS/MTS would be slightly lower than expected, provided that the feed conditions are the same. The temperature trend with different steam:dry gas ratios, as simulated in a simulation tool, is shown in FIG. 7. The maximum bed temperature at an S:G ratio of 0.493 (normal case) is 421°C; and, at S:G ratios of 0.6 and 0.75, the maximum bed temperatures are 419°C and 416°C, respectively. The featured example shows this simple method by which boiler leaks can be detected before it is too late. Frequent gas sample analysis and mass balance are the keys in detecting any boiler feed water ingress into the shift reactor. REFERENCES Broadhurst P. V. and P. E. J. Abbott, “Improving hydrogen plant performance, Part 2,” PTQ, 2002. 2 Anderson R., P. V. Broadhurst, D. Cairns, F. E. Lynch and C. Park, “The chemistry within your catalysts, Part 3—Water gas shift and methanation,” Nitrogen & Syngas 310, March–April 2011. 3 Lynch, F., and S. Appleton, “Water and your shift converter—Hero or villain?,” Ammonia Technical Manual 2003, Proceedings of AIChE Ammonia Plant Safety Symposium, 2003 1

K. R. RAMAKUMAR is a technical service engineer for Johnson Matthey and is based in Dubai, UAE. He has more than 10 years of process, operations and commissioning experience in hydrogen and hydrocracking plants. Mr. Ramakumar holds a BS degree in chemical engineering.

Fluid Flow and Rotating Equipment H. PANDYA, Saipem India Projects Pvt. Ltd., New Delhi, India; and A. M. FANTOLINI, Saipem, Milan, Italy

Limit the rate of change of fuel gas properties with mixing drum The fuel gas mixing drum is an integral part of many plants, especially LNG plants. Any change in fuel gas composition, properties or pressure can lead to operational problems in fuel-using systems, such as gas-fired turbines. These systems are designed to accept a certain degree of change of fuel gas properties; however, their burners are sensitive to the rate of this change. To limit this, fuel gas mixing drums are typically used. FIG. 1A shows a typical sketch of a fuel gas mixing drum. It consists of two separate sections, with a knockout drum (KOD) in the bottom section and a central riser pipe with a number of baffles in the top section. Fuel gas enters from the bottom section and flows to the top section, or vice versa. The riser contains a series of holes in each baffle section, so gas from the riser holes mixes with gas from the baffle sections. Baffles are sloped to allow the free draining of liquid. This arrangement ensures proper mixing and gradual changes in any process parameters. The KOD section (FIG. 1) is designed as a normal two-phase separator, such as a compressor suction drum. Here, the mixing drum design procedure is reviewed for operational optimization. Number of mixing stages in top section. The fuel gas mixing drum en-

sures gradual change in fuel gas properties, such as the Wobbe Index, and also acts as a buffer vessel to limit pressure variation. The Wobbe Index is an indicator of the interchangeability of fuel gases. Some gas turbine vendors use

different terminologies for this parameter, such as the Modified Wobbe Index and Gas Index, and to include temperature as an additional variable; however, these are all derivatives of the Wobbe Index. In the following discussion, the equipment is designed to limit the rate of change of the Modified Wobbe Index (MWI), defined as: MWI =

LHV T.SG

(1)

If the feed gas composition changes, i.e., from heavier to lighter, the MWI at the outlet of the drum will change gradually from value MWIOLD (heavy gas) to value MWINEW (light gas). ΔMWI =

( MWI New − MWIOld ) MWIOld

FIG. 1B. Assuming that the total number of

compartments are n, for the nth compartment, if the change in fuel gas property starts to happen at time t = t + Δt, then the property balance for the nth compartment at time t = t + Δt can be written as: ⎛ F ⎞ MWI n,t+Δt = MWI n,t + MWI n−1,t ⎜ n−1 ⎟ ⎝ Vn ⎠ ⎛ F /n ⎞ Δt + MWI New⎜ ⎟ Δt − (4) ⎝ Vn ⎠ ⎛ F ⎞ MWI n,t ⎜ n ⎟Δt ⎝ Vn ⎠ Gas outlet Mixing drum

×100 (2)

If the maximum allowable rate of change of the MWI is R Allowable [%/sec] as specified by the gas turbine vendor, the residence time in the mixing drum required to meet the rate change requirement will be: ΔMWI ≤ R Allowable ⎛ VInitial ⎞ (3) ⎜ ⎟ ⎝ F ⎠ Volume, VInitial , is the initial volume of the mixing drum required to meet design parameters, if no internals, such as riser baffles, are present. Due to the presence of internals, mixing is increased and the required volume is consequently decreased. To relate the required vessel volume with the number of compartments, an analytical model of equipment is shown in

Riser

Baffle

Drain holes

Feed gas in Knockout drum Liquid outlet

FIG. 1A. Typical sketch of a fuel gas mixing drum. Hydrocarbon Processing | MARCH 2016 67

Fluid Flow and Rotating Equipment Considering the maximum variation is at the beginning of the disturbance, to calculate the initial tangent of the variation curve for the last compartment, the differential of Eq. 4 with respect to time is taken.

Fn=F (F/n)

F(n-1) Vn-1

V4

(F/n)

F(n-2)

(F/n)

⎛ F ⎞ ⎛ d ( MWI ) ⎞ = MWI n−1,t ⎜ n−1 ⎟ + ⎜ ⎟ ⎝ dt ⎠n,t+Δt ⎝ Vn ⎠

F3 (F/n)

F2 V2

Vn

(F/n)

⎛ F /n ⎞ ⎛ F ⎞ MWI New⎜ ⎟ − MWI n,t ⎜ n ⎟ ⎝ Vn ⎠ ⎝ Vn ⎠

V3

(F/n)

V1

(5)

⎛ F − ( F /n ) ⎞ ⎛ d ( MWI ) ⎞ = MWI Old ⎜ n ⎟ ⎜ ⎟ ⎝ dt ⎠n,t+Δt Vn ⎝ ⎠

F

⎛ F /n ⎞ ⎛ F ⎞ + MWI 'New ⎜ ⎟ − MWIOld⎜ n ⎟ (6) ⎝ Vn ⎠ ⎝ Vn ⎠

FIG. 1B. An analytical model of a fuel gas mixing drum. 1.0

Assumptions: • Flow through riser in each compartment is equal (F/n). • Inlet flow (F) to vessel does not change with time. • All compartments have equal residence time (Vn/Fn = constant). • Area of individual hole is small compared to the cross-sectional area of the riser pipe. • Turbulent flow of gas in riser up to last section of riser (i.e., Reynolds number > 5,000).

Ratio, VFinal/VInitial

0.9 0.8 0.7 0.6 0.5 0

5

10

15

20 25 30 35 Number of compartments

40

45

50

FIG. 2. The number of compartments vs. volume ratio.

Example 1 In a plant, fuel gas is changed from a light gas (MWI = 44 MJ/m3) to a heavier gas (50.2 MJ/m3). Fuel gas flow is 0.5 m3/sec. Design a fuel gas mixing drum top section to ensure maximum rate of change in MWI of 0.3 %/sec. Solution: Change in MWI = (50.2-44)/44 × 100 = 14.1% VInitial From Eq. 3, 23.5 m3 Number of compartments assumed: 7 Riser diameter = 12 in. = 0.3 m VFinal: (7 + 1)/(2 × 7) × 23.5 = 13.4 m3 Assume for vessel: Length/diameter ratio: 2.2 π/4 × D2 × (2.2 × D) = 13.4 D = 1.7 m, assumed D = 1.7 + 0.3 = 2.0 m L = 2.2 × 2.0 = 4.4 m Length of first compartment, from Eq. 10 L1 = 4.4/(7 × (7 + 1)/ 2) = 157 mm However, to enable a manhole at first and last compartment, L1 assumed = 800 mm Hence, overall length of top section = 5.1 m. Bottom KOD length = 3.6 m (from standard 2-phase vessel calculation) Height of bottom head (elliptical assumed) = D/4 = 0.5 m Total height = 9.2 m 68 MARCH 2016 | HydrocarbonProcessing.com

⎛ F /n ⎞ ⎜ ⎟ ⎝ Vn ⎠

(7)

where Vn can be written as:

At the start of variation at time t = t, MWIn–1,t = MWIn,t = MWIOld :

F1

⎛ d ( MWI ) ⎞ = ( MWI New − MWI Old ) ⎜ ⎟ ⎝ dt ⎠n,t+Δt

V1+V2+V3+………+Vn = VFinal

(8)

V2 = 2V1 , V3 = 3V1 , ………Vn = nV1 V1+2V1+3V1+................+nV1 = VFinal (9)

⎛ n (n+1) ⎞ VFinal → V1 = (10) ⎜ ⎟V =V ⎛ n (n+1) ⎞ ⎝ 2 ⎠ 1 Final ⎜ ⎟ ⎝ 2 ⎠ VFinal VFinal Vn= n = (11) ⎛ n (n+1) ⎞ ⎛ (n+1) ⎞ ⎜ ⎟ ⎜ ⎟ ⎝ 2 ⎠ ⎝ 2 ⎠

From Eqs. 7 and 11 and dividing by MWIOld: ⎛ 1 ⎞ ⎛ d ( MWI ) ⎞ ⎜ ⎟= ⎜ ⎟ ⎝ dt ⎠n,t+Δt ⎝ MWIOld ⎠

(12)

⎛ ⎞ ⎜ ⎟ ⎜ ( MWI New − MWIOld ) ⎜ F /n ⎟⎟ ⎜ VFinal ⎟ MWIOld ⎜ ⎛ (n+1) ⎞ ⎟ ⎟ ⎟ ⎜ ⎜ ⎝ ⎝ 2 ⎠ ⎠ ⎛ 1 ⎞ ⎛ d ( MWI ) ⎞ ⎜ ⎟= ⎜ ⎟ ⎝ dt ⎠n,t+Δt ⎝ MWIOld ⎠ ΔMWI ⎛ ⎛ 2n ⎞ ⎞ ⎟ ⎟ ⎜ VFinal ⎜⎝ n+1 ⎠ ⎟ ⎜ F ⎜ ⎟ ⎝ ⎠

(13)

Comparing Eq. 3 with Eq. 13, ⎛ V ⎞ ⎛ n+1 ⎞ ⎛ 2n ⎞ VInitial = ⎜ ⎟VFinal → ⎜ Final ⎟ = ⎜ ⎟ (14) ⎝ n+1 ⎠ ⎝ VInitial ⎠ ⎝ 2n ⎠

Eq. 14 shows that VFinal depends upon the number of compartments. For n = 1, the ratio of VFinal to VInitial is equal to 1. However for n = infinite, it is 0.5, which infers that in no case can VFinal be less than 50% of VInitial. FIG. 2 shows that the ratio initially falls steeply up to n = 5 and then falls gradually. Hence, it is preferred to have the number of compartments be between five and 10. The calculated volume excludes riser volume; therefore, the diameter of the vessel must be increased to adjust for riser volume. Increased diameter can be calculated as the square root of the sum of the square of vessel diameter and riser diameter. As a

Fluid Flow and Rotating Equipment conservative approach, the riser diameter is often simply added to the vessel diameter. Computational fluid dynamics simulations are used to support and verify the sizing of such critical items, leading to the optimization of the mixing inside the drum (FIG. 3), the identification of the trajectories of flow from the holes, and the definition of the velocity vectors in different points of the drum. Pressure drop calculation. The mixing drum pressure drop includes the inlet/ outlet nozzle pressure drop, KOD pressure drop, baffle section pressure drop, riser pipe and holes pressure drop. Inlet/outlet nozzle pressure drop and KOD pressure drop are calculated as standard vessel pressure drop. Baffle section pressure drop is negligible and so can be ignored. The combined total of all these pressure drops are in the range of 5 kPa–10 kPa (excluding riser pipe and holes pressure drop). Finally, the riser diameter and the diameter of holes can be selected through using iterative calculations. At first, the riser diameter is set the same as the feed gas inlet diameter, and the holes size is set at 30 mm. The number of holes are adjusted based on mechanical considerations: holes mus be placed for turbulent conditions at the exit while maintaining mechanical resistance in the riser, and to satisfy the pressure drop requirement. If the velocity is sufficiently high (i.e., turbulent conditions are reached in the hole), and the number of holes is still low, the first trial calculation may be acceptable. If the velocity is low and no turbulent conditions are observed at the exit (max. allowable 100 m/s), or the number of holes is too high, a recalculation has to be made with a different number of holes to verify the velocity, the Reynolds number and pressure drops across the holes. Riser pressure drop calculation. Pressure drop in a perforated pipe distributor for turbulent flow with roughly uniform flow distribution:1 ⎛ 4 fL ⎞ ρv 2 Δp = ⎜ − 2K ⎟ i ⎝ 3d ⎠ 2 ⎛ 4 fL ⎞ if ⎜ ⎟