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NOVEMBER 2010

HPIMPACT

SPECIALREPORT

BONUSREPORT

LCFS impact oil sands refiners

PLANT SAFETY AND ENVIRONMENT

HEAT TRANSFER

US demand for lubricants increasing

Methods to mitigate CO2 and other air emissions

New designs foster efficiency and uptime

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NOVEMBER 2010 • VOL. 89 NO. 11 www.HydrocarbonProcessing.com

SPECIAL REPORT: PLANT SAFETY AND ENVIRONMENT

31

Stretch in technology and gaps in process safety for the hydrocarbon industry Industry faces challenges to prevent and control hazards M. S. Mannan

33

Better water management: A crucial and growing requirement Technologies for water reuse and recycling needed now G. Messina

35

Water among causes for storage tank explosion

Cover The Kårstø gas processing plant on the west-coast of Norway is the largest of its type in Europe. The installation plays a key role in the transportation and treatment of gas and condensate (light oil) from important areas on the Norwegian continental shelf. Thirty fields are linked to Kårstø via pipelines. Every day millions of cubic meters of gas and non-stabilized condensate flow into the plant, where the heavier components are removed by separation. The remainder, known as dry or sales gas, is exported by pipeline to the Continent.

Reinvestigation uncovers true accident events M. Ferjencik and B. Janovsky

41

Consider new analysis for flares

49

Designing the correct pressure-relieving system

55

Customize operator training for your thermal oxidizers

61

Applying dynamic models in designing safety systems can reduce capital costs Z. Urban, M. Matzopoulos, J. Marriott and B. Marshall

Use these relief rate calculations for gas thermal expansion as a cause for overpressure S. Rahimi Mofrad

This case history shows the benefits of site-specific programs in new equipment installations T. Gilder, D. Campbell, T. Robertson and C. Baukal

Emergency response planning— start at the plant design stage Follow these guidelines for a safer facility R. Saini

65

Optimized fired heater control

69

Consider real-gas modeling for turboexpanders

73

Predictive emissions monitoring helps reduce stack air emissions

New visualization methods expose problems with traditional designs K. Kaupert

Consider switching to Internet protocol surveillance Here’s a checklist to make the jump M. S. Wilson

ENGINEERING CASE HISTORIES Case 59: Heat-up rates and thermal cracking

81

17 LCFS will adversely affect oil-sands crude refiners 17 US demand for lubricants to reach 2.25 billion gallons in 2014 17 Linde supplying hydrogen technology to US BMW plant 18 Oil and gas reserves increased 3 percent in 2009

Residual oxygen measurement principle lowers emissions and improves efficiency A. J. Mouris

New technology reduces compliance costs while optimizing operations R. Hovan

79

HPIMPACT

A good analysis is usually better than speculation T. Sofronas

DEPARTMENTS 7 HPIN BRIEF • 23 HPINNOVATIONS • 27 HPIN CONSTRUCTION 29 HPI CONSTRUCTION BOXSCORE UPDATE 82 HPI MARKETPLACE • 85 ADVERTISER INDEX

COLUMNS 9 HPIN RELIABILITY Wire mesh vs. wire size for temporary strainers 11 HPIN EUROPE Consultants stand by need for swingeing cuts 13 HPINTEGRATION STRATEGIES Future of the collaborative process automation system 15 HPI VIEWPOINT Weighing on dualtemperature control 86 HPIN WATER MANAGEMENT Utility water boot camp for process engineers—Part 3


Italian design A masterpiece

www.HydrocarbonProcessing.com Houston Office: 2 Greenway Plaza, Suite 1020, Houston, Texas, 77046 USA Mailing Address: P. O. Box 2608, Houston, Texas 77252-2608, USA Phone: +1 (713) 529-4301, Fax: +1 (713) 520-4433 E-mail: editorial@HydrocarbonProcessing.com www.HydrocarbonProcessing.com Publisher Bill Wageneck bill.wageneck@gulfpub.com EDITORIAL Executive Editor Stephany Romanow Process Editor Tricia Crossey Reliability/Equipment Editor Heinz P. Bloch News Editor Billy Thinnes European Editor Tim Lloyd Wright Contributing Editor Loraine A. Huchler Contributing Editor William M. Goble Contributing Editor Y. Zak Friedman Contributing Editor ARC Advisory Group (various) MAGAZINE PRODUCTION Director—Editorial Production Sheryl Stone Manager—Editorial Production Angela Bathe Artist/Illustrator David Weeks Manager—Advertising Production Cheryl Willis ADVERTISING SALES See Sales Offices page 84. CIRCULATION +1 (713) 520-4440 Director—Circulation Suzanne McGehee E-mail: circulation@gulfpub.com SUBSCRIPTIONS

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If you would like to have a recent article reprinted for an upcoming conference or for use as a marketing tool, contact Foster Printing Company for a price quote. Articles are reprinted on quality stock with advertisements removed; options are available for covers and turnaround times. Our minimum order is a quantity of 100. For more information about article reprints, call Rhonda Brown with Foster Printing Company at +1 (866) 879-9144 ext 194 or e-mail rhondab@FosterPrinting.com. HYDROCARBON PROCESSING (ISSN 0018-8190) is published monthly by Gulf Publishing Co., 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 © 2010 by Gulf Publishing Co. All rights reserved. Permission is granted by the copyright owner to libraries and others registered with the Copyright Clearance Center (CCC) to photocopy any articles herein for the base fee of $3 per copy per page. Payment should be sent directly to the CCC, 21 Congress St., Salem, Mass. 01970. Copying for other than personal or internal reference use without express permission is prohibited. Requests for special permission or bulk orders should be addressed to the Editor. ISSN 0018-8190/01. www.HydrocarbonProcessing.com

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HPIN BRIEF BILLY THINNES, NEWS EDITOR

BT@HydrocarbonProcessing.com

The US Environmental Protection Agency (EPA) has waived a limitation on selling fuel that is more than 10% ethanol for model year 2007 and newer cars and light trucks. The waiver applies to fuel that contains up to 15% ethanol—known as E15—and only to model year 2007 and newer cars and light trucks. This represents the first of a number of actions that are needed from federal, state and industry towards commercialization of E15 gasoline blends. A decision on the use of E15 in model year 2001 to 2006 vehicles will be made after EPA receives the results of additional Department of Energy testing, which is expected to be completed in November. However, no waiver is being granted this year for E15 use in model year 2000 and older cars and light trucks—or in any motorcycles, heavy-duty vehicles or non-road engines—because currently there is no testing data to support such a waiver. Since 1979, up to 10% ethanol or E10 has been used for all conventional cars and light trucks, and non-road vehicles.

GE has snapped up Dresser for a cool $3 billion. The deal includes all of the Dresser businesses that provide products and services for compression, flow technology, measurement and distribution infrastructure for customers in more than 150 countries. Dresser’s extensive, global-installed base of products generate aftermarket service revenues in excess of 40% of total revenues. Morgan Stanley acted as the exclusive financial advisor to Dresser on this transaction.

Strong Petrochemical Holdings Ltd. has entered into an exclusive marketing agreement with Eurocontrol Technics Inc. Eurocontrol will appoint Strong as the exclusive sole agent to market and sell Eurocontol’s Petromark fuel authentication system to potential customers in China, Hong Kong and other defined target markets throughout Asia. This exclusive agreement is contracted for an initial period of three years and is renewable for further periods of three years. Strong will act as the exclusive sole agent, responsible for the marketing and sale of the products in the territories and Global Fluids International S.A., a wholly owned subsidiary of Eurocontrol, will be responsible for the design and technical support to the consortium. Strong will bring to the consortium its extensive business connections in the target markets with participants in the petrochemical complex.

Petrofac has a contract with the government of Sharjah, UAE, to take over operational responsibility and facilities management of the Sajaa gas plant and related assets, located approximately 30 km from Sharjah, UAE. The five-year contract was awarded following a competitive open bidding process and is worth in excess of $250 million. The government of Sharjah, acting through the Sharjah Petroleum Council, holds a 60% participating interest in the Sharjah gas and associated liquids concession.

Marathon Petroleum Co. LP (MPC) is selling most of its downstream assets in Minnesota to ACON Investments, LLC, and TPG Capital, LP. ACON and TPG formed Northern Tier Energy LLC to operate the assets as a stand-alone company. Included in the transaction will be the 74,000 bpd St. Paul Park refinery and associated terminals, 166 SuperAmerica convenience stores (including six stores in Wisconsin), SuperMom’s LLC, SuperAmerica Franchising LLC, interests in pipeline assets in Minnesota and associated inventories. The total sales value is approximately $900 million, including Northern Tier preferred stock with a stated value of $80 million. Approximately $300 million of the total sales value is for the inventories associated with these operations. The agreement also contains earnout and margin support components where Marathon could receive up to an additional $125 million over eight years or may be required to provide up to $60 million of margin support to the buyers, subject to certain conditions. HP

■ Better gas storage via MOFs Natural gas-powered vehicles may soon be able to travel double the distance on a single tank—due to metal organic frameworks (MOFs). BASF research scientists have developed an innovative method for the solventfree industrial-scale manufacture of those materials for better gas storage. MOFs produced by the new method are currently being trialed for natural gas storage in heavy-duty vehicles. With their special structure and large surface area, MOFs open up new opportunities for alternative propulsion systems, in catalysis, as nanoreactors, and in drug delivery, making them hugely interesting both for industry and university research. “This substance class opens up new areas of applications in material science. We are delighted at this significant advance in industrialscale production, which is a crucial requirement for the commercial use of these fascinating materials,” said Dr. Friedrich Seitz, head of research chemicals at BASF. BASF has been working toward industrial-scale synthesis of metalorganic frameworks for the past 10 years. MOFs are highly crystalline structures with nanometer-sized pores that allow them to store hydrogen and other high-energy gases. The larger specific surface area and high porosity on the nanometer scale enable MOFs to hold relatively large amounts of these gases. The pores are adjustable in terms of size and polarity and so can be fine-tuned for specific applications. Used as storage materials in the natural gas tanks of municipal utility vehicles, MOFs offer a docking area for gas molecules, which can be stored in higher densities as a result. The larger gas quantity in the tank increases the vehicle’s range. An advantage of the production method developed by BASF is that it uses no organic solvents. The simple method gives a higher material yield from an aqueous medium and is suitable for existing BASF production plants. MOFs were discovered toward the end of the 1990s by US chemist Omar M. Yaghi. HP HYDROCARBON PROCESSING NOVEMBER 2010

I7


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

Wire mesh vs. wire size for temporary strainers Even an experienced reliability professional may find some customary terminology dealing with wire gauge and wire mesh to be confusing. Our HPIn Reliability column of April 2009 mentioned pump suction strainers made of 20-gauge wire mesh and a few words about wire gauge will be helpful. Wire gauge (also known under the alternative spelling “wire gage”) is indicative of electrical wire diameter, with primary implications for allowable electric-current flow (amperes). According to working tables published for Standard Annealed Solid Copper Wire,1 20-gauge wire has a diameter of 0.032 in. and, at about 40°F, has a resistance of approximately 11 ohms per 1,000 ft of length. While this month’s column has nothing to do with electric current, we want to assume that the 20-gauge stainless-steel wire we would like to see inside the three-mesh guard screen would be woven from wire that has a diameter of 0.032 in. The terms wire mesh and wire cloth are often used interchangeably. They denote a metal wire weave with its various implications regarding available flow-through area as a percentage of the total square inches of wire cloth area. Our April 2009 column meant to recommend placing a wire mesh made from 20-gauge wire inside a three-mesh guard screen. So, to clarify: While wire mesh made from 1⁄32-in. wire is to be placed inside a three-mesh guard screen, no 20-mesh wire cloth is involved. The three-mesh guard screen is usually recommended for temporary strainers to reinforce a finer mesh. • “Three-mesh” means wire cloth (or guard screen in the illustration published in April 2009). On a piece of paper, that would be three wires per in. in the x-direction (left-to-right) and three wires-per-in. in the y-direction (up-and-down). Similarly, eight mesh would be eight wires in the x-direction and eight wires in the y-direction; or 200 mesh would be 200 wires in the x-direction and 200 wires in the y-direction, etc. • In a three-mesh screen there are, therefore, 3 x 3 = 9 openings. • In a three-mesh screen, the distance from the center of one wire to the center of an adjacent wire is 1⁄3 in., or 0.333 in. • One manufacturer2 of wire cloth offers three-mesh cloth made with wires ranging from 0.031 in. to 0.162 in.—in each case, there would be nine openings and the distance from the center of a wire to the center of an adjacent wire would always be 0.333 in. • If one were to pick the 0.162-in. thick wire, one would have greater strength than if one picked the 0.031-in. wire. • Likewise, if one picked the 0.162-in. wire diameter, the resulting open area would only be 26%; whereas, if one selected the 0.03-in. wire diameter, there would be an 82% open area. • For temporary strainers in process pump piping, one would pick a three-mesh guard screen with a wire diameter of about 0.135 in. To reiterate, the wire-to-wire center distance would be 0.333 in. and there would be nine openings. The wire-cloth manufacturer’s tables (obtained from the Internet2) for this diameter give an opening width of 0.198 in. (remember that opening width plus

wire diameter equals 0.333 in.). The tables also tell us that we have ~35% of the area open for liquid flow and ~65% of the area would be taken up by the 0.135-in. diameter guard screen wires. Next, one would determine the wire mesh utilizing the 20-gauge wires. That particular wire cloth, to be placed inside the three-mesh guard screen, should use 20-gauge wire. Reference 2 shows the related wire diameter (0.032 in.) in the middle of a six-mesh table. The openings are 0.1347 in., and adding the two numbers gives us 0.1667 in.—six squares per in. Using the above information and Fig. 1 will equip us to ask a draftsman to design a pointed cone—the three-mesh guard screen—with a surface area about three times that of the pipe or spool piece cross-sectional area. The six-mesh metal cloth would go on the inside of the cone and only then could we call it a strainer. On any permanent strainer installation, it would be wise to monitor the delta-p (the pressure drop) between the mesh upstream and downstream sides. Clogged strainers can cause serious machine malfunction and costly damage. HP 1 2

LITERATURE CITED Marks’ Standard Handbook for Mechanical Engineers, 7th Edition (1969), McGraw-Hill Book Company, New York, New York. www.mcnichols.com/products/wiremesh.

The author is Hydrocarbon Processing’s Equipment/Reliability Editor. A practicing consulting engineer with close to 50 years of applicable experience, he advises process plants worldwide on failure analysis, reliability improvement and maintenance cost-avoidance topics. He has authored or coauthored 17 textbooks on machinery reliability improvement and over 470 papers or articles.

Identification tab at top for raised-face flanges. Mount on flange bolts. Preferred flow direction

Strainer screen Spool piece Screen seam

Strainer assembly

Mount mesh made from 20-gauge wire inside a three-mesh guard screen and stagger the longitudinal seams.

Install gasket on each side of strainer flange.

12-gauge x 1-in. wide identification tab.

½-in. 12-gauge ASTM “A” 167 type 316

½-in.

90°

Screen section See detail “A”

FIG. 1

1-in.

Detail “A”

Both sides of strainer flange must be free of gouges, weld spatter or other imperfections that might impair proper gasket sealing.

Temporary strainer for pump suction piping.

HYDROCARBON PROCESSING NOVEMBER 2010

I9


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

Consultants stand by need for massive capacity cuts In November 2009, I wrote about how the world’s downstream consultants seemed almost to be vying to ring up the most pessimistic prognosis for refinery closures. A year on, there’s little to suggest they were off the mark. Talking to one of them today, they’re just as bearish as they were then. You may recall that JBC Energy in Vienna told me last year that 2.8 million bpd (MMbpd) of European refinery capacity must close. By February of this year, they increased that estimate by nearly 600,000 bd. How much to cut? From its work for OPEC, the US consult-

ing group, Ensys said that 10 MM bpd of cuts would be required to bring global utilization rates up to 85% by 2020. Energy Market Consultants (EMC) in London, part of Facts Global Energy, had perhaps the most dire warning, putting the level of necessary reductions at 11 MMbpd by 2020. This was when I asked Leif Nilsson, the canny trading manager at Preem, by then specializing in storing up oil for better times ahead, for his advice for refiners. “Margins are going to be bleak,” he said. “Just try and sit it out.” Margins are still bleak. Margins have come back some

since the depths of the economic crash last year, but they’re still way below the pre-crash glory days. EMC’s variable cost FCC margin is hovering around $1.50/bbl, and, with some cuts of the barrel, it is heavily negative. Refiners have clearly concluded that they need to take action to reduce their exposure. Shell has announced the closure of one site, albeit at the far extent of the Atlantic Basin in Montreal. It has successfully concluded the sale of a German site, Heide, to a private investor, and is said now to be in exclusive talks to sell its Gothenburg, Sweden refinery to a Finnish company with fuel retailing interests. In its most recent results presentation to analysts, Total reiterated its plans to reduce its refining output by 500,000 bpd by 2011. There was, however, no new information on its long-standing effort to sell the UK Lindsay refinery, or its UK retail network. Meanwhile, Total’s abortive attempt to negotiate redundancies and the closure of its Dunkirk, France, refinery are thought to be what has given some potential investors in Europe’s downstream industry pause for thought. I hear that Essar of India’s team, which has been looking over key assets like the 246,000-bpd Shell Stanlow refinery, got itchy feet after watching, as on June 30, a French court ordered Total to restart its Dunkirk refinery within 15 days, or face a €100,000 fine. Hope floats. There are some small signs of light at the end of this dark tunnel. For one thing, a key distillate market is in flux. Heating oil futures have been in contango since the massive

demand destruction that accompanied the 2008 money supply crisis. With no buyers now, futures markets have consistently reflected premiums for oil delivered in the future. It’s a bearish, rather than optimistic, signal, suggesting a belief that things can only get better because they’re so bad now. In just the last few days, as I write, that structure has flipped around, although just for the coming months before winter. Suddenly and admittedly compared to a very low baseline, Europe is the top destination for surplus cargoes of distillate. A surge of opportunistic vessel fixtures this week indicated that US and Asian traders and refiners, as predicted in my October column incidentally, are attempting to capture some of a $17/metric ton spread between European ULSD/50-ppm heating oil and spot prices in the US. Reflecting the improved demand, the OPIS Spot Northwest Europe Gasoil Cargoes to Brent Crude crack averaged $11.24/ bbl for the first three weeks of September this year, compared to $5.74/bbl for the same three weeks a year earlier. Still too much spare capacity. So when I spoke to Roy Jordan at EMC this week, I asked if he thought the reversed market structure was indeed a sign that things are on the up for European refiners. “Unfortunately, you’ve got to see it in the context of a structural surplus of gasoline and an inability to export to the US,” he said. “Yes, that heating oil crack is improved, but you have to see it against very weak fuel, gasoline and naphtha cracks.” Jordan says that the heating oil market, itself driven to a great extent by diesel demand, is Europe’s essential oil. Take that profitable cut away from European refiners, and they might as well shut up shop altogether, he suggests. EMC’s view of the recent backwardation? “We’ve got it down to the refinery maintenance season,” he says. And so, as EMC prepares new figures for a forthcoming refining outlook, I asked Jordan if EMC is sticking to its guns on the need for cuts. When I last spoke to EMC, they were estimate a need for 7 MMbpd of reductions by 2015, with a further 4 MMbpd necessary by 2020. “There’s no retreat since then,” says Jordan. “We’re more confirmed than ever. The spare capacity is enormous and margins are not recovering.” HP

The author is HP’s European Editor and also a specialist in European distillate markets. He has been active as a reporter and conference chair in the European downstream industry since 1997, before which he was a feature writer and reporter for the UK broadsheet press and BBC radio. Mr. Wright lives in Sweden and is the founder of a local climate and sustainability initiative. HYDROCARBON PROCESSING NOVEMBER 2010

I 11


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HPINTEGRATION STRATEGIES DAVE WOLL AND LARRY O’BRIEN, CONTRIBUTING EDITORS dwoll@arcwb.com

Future of the collaborative process automation system ARC conceived the Collaborative Process Automation Systems (CPAS) in 2002 in response to DCS end users’ requests for us to provide a vision of how process automation systems should evolve. Clients and others continue to ask ARC where the automation market is headed. The DCS came on the scene in 1975 and smart field devices arrived about five years later. Since then, ARC has observed incremental innovation, but nothing revolutionary. With this in mind, we recast the request for ARC’s view into the future as an effort to identify innovation we think the market needs to consider in the near future.* Smart field devices are not smart enough. At present, 50% to 75% of field-device downtime is caused by lack of confidence in the measurements. The current generation of smart field devices is capable of communicating basic device diagnostics along with the measurements. While this has been acceptable in many industries, we would also like to see quantitative measurements of the relative health of a device and, if there is an issue, how long the device will continue to operate dependably. When will distributed control become truly distributed? Manufacturing challenges are changing, and automation

must change to stay in control. We now deal with a high degree of uncertainty in an increasingly dynamic environment. Manufacturing assets continue to become more costly to purchase. Expansions to existing manufacturing assets are expensive, and existing assets almost impossible to replace. While automation can be viewed as a solution to a certain degree, these systems could provide more. What have been dubbed distributed control systems in the past are really not all that distributed, since, to a large degree, they still tie functional requirements to architectural constraints. Future requirements need intelligent agents. Manu-

facturing challenges continue to change, and the current IEC 61131-3 standard will not be able to satisfy the execution semantics and new requirements for a truly distributed and flexible automation system. The Technical Committee 65 received a new work proposal (NWP) to standardize application of function block software modules in distributed industrial-process measurement and control systems. The resulting standard, IEC 61499, will facilitate the manufacturing agents that will shape the next-generation automation systems. Multi-vendor control. The technology exists (or is certainly

feasible) to make the basic control system a multi-supplier structure. For example, if the control system LAN were based on FOUNDATION fieldbus HSE, then just like field instruments can come from different suppliers, the process controllers could also come from different suppliers.

Online version upgrade. Process automation system supplier has at least one major version release every year. Unlike a maintenance release, a version release usually requires the system to be taken out of service (thus interrupting production). With most of today’s systems, it’s possible to perform a twostep version “hot” upgrade without shutting down by loading the new version into the redundant side, then committing the new software to the primary side of the system and the process. We feel that this is unreasonably risky and only marginally practical. We feel a three-step approach where the changes can be validated is reasonable. Business-to-operations integration. Normally, integrat-

ing two different systems, especially systems as diverse as a business and manufacturing system, can be complex and difficult. This is where Reference Standards can make a major contribution. Their primary value is that they contain well-understood and well-documented work processes that lead to handing off proper information. Several years ago, several senior practitioners on the business and manufacturing sides recognized the value of integrating Supply Chain Management Consortium reference model (SCOR) with the manufacturing reference model, ISA-95, and have began to address the problem. Application executive of the future. The application

executive, a core CPAS function, monitors the health of applications running in a collaborative and synchronized manner with other applications, the system itself, and the personnel using the system. Today, this is accomplished largely in an environment combining proprietary structures, de facto standards and enabling technologies. One potential scenario for the future of the application environment is to have this completely based on standards such that no single supplier has to “own” the environment or platform for it to function properly. In this future vision, the user would purchase applications from the supplier that provided the best fit for the required functions. It would be loaded into the system and would auto-configure to adapt to the operating system, communications protocols, and whatever hardware requirements are present. HP *Note:

Readers can visit www.arcweb.com/res/cpas for more information on ARC’s latest research into CPAS.

Dave is a vice president of consultingconsulting at the ARC Advisory Group. He has Larry Woll O’Brien is part of the automation team at ARC covering the been associated with and applyingeditor. process forfor over 35 years. process industries, anddefining an HP contributing He automation is responsible tracking the Mr. Wollfor is process currentlyautomation focused onsystems assisting major users their market (PASs) and has developing authored the PASprocess market autostudmation system strategies. ies for ARC since 1998. Mr. O’Brien has also authored many other market research, strategy and custom research reports on topics including process fieldbus, collaborative partnerships, total automation market trends and others. He has at been with ARC since Larry O’Brien is part of the automation consulting team ARC covering the January and started his career with market research in the instrumentation process 1993, industries. He is responsible for tracking the market forfield process automation markets. systems (PASs) and has authored the PAS market studies for ARC since 1998. HYDROCARBON PROCESSING NOVEMBER 2010

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HPI VIEWPOINT ALLAN KERN, GUEST COLUMNIST Allan.Kern@yahoo.com

Weighing on dual-temperature control I’d like to weigh in again on the continuing debate regarding distillation column control.1–4 Attention has now turned to the practicality of “dual temperature” control,which is the simultaneous control of temperature on both the overhead and bottoms of a distillation column, typically with top temperature cascaded to reflux flow and bottom temperature cascaded to reboiler heat input. My short answer agrees with Zak Friedman: Beware, all ye who wander here! Attempting dual-temperature control is most often undertaken by those oblivious to its perils. Industry is littered with its wreckage in the form of columns configured for dualtemperature control, but with one of the controls abandoned— either the reflux or reboiler is operated in automatic or manual mode, not cascaded. This is unfortunate from both composition control and energy conservation points of view. The lesson from all the wreckage (and from this debate) is that successful dual-temperature control takes more preparation than most casual control strategy reconfigurations get. Appropriate preparation in this case would include a rigorous simulation of the several variables that affect overall separation because dual-temperature control really means controlling one temperature plus overall separation (think Δ-temperature between the top and bottom). For practical purposes, separation is usually considered constant. When Mr. Shinskey or Mr. Friedman points out that a onedegree change in bottom temperature will result in approximately a one-degree change in the overhead (Δ-temperature remains the same), this is another way of saying the same thing. If separation is constant, it cannot be controlled, and therein is the problem that many people fail to realize, at least initially. To the extent that overall separation does vary, it is usually due to uncontrollable factors, such as feedrate and quality, rather than controllable ones. That said, if you have the luxury of controlling feed or a feed analyzer, responsive pressure control, which affects vapor velocity and tray efficiency, ) excess stages, a reliable simulation, good operators and economic or equipment limitations that prohibit doing the split in two columns, then perhaps dual-temperature control is worth a try. That’s my long answer, which more agrees with Shinskey, and which can be shortened to: if it must be done, it might be possible! A third answer worth mentioning is that the traditional solution (or compromise) to this dilemma is to put one temperature control on the top or bottom (whichever purity is the higher priority or harder to achieve) and operate the other end of the column on ratio control. This results in good stability, depending on feed and other variables, and keeps energy consumption in check. The ratio is usually reboiler heat-to-feed (when top temperature is controlled) or reflux-to-feed (when bottom temperature is controlled). The latter option (reflux ratio) has a couple “material balance” risks. If all the light material is not removed with the distillate, it can accumulate in the overhead and lead to erratic pressure and temperature behavior. It’s possible to run the reflux drum empty.

There are no such limitations with the former option (reboiler ratio), and indeed, it is probably the single most common (successful) column control configuration in the industry. It’s not as good as the grail of stable dual-temperature control, and is a good measure better than running the reflux or reboiler at a fixed setpoint, both in terms of constant separation and energy conservation. Yet another answer is special cases. On columns that have side-draws, additional boil-up does not ultimately return as reflux, so dual-temperature control can work well. This is well known on main fractionators that often have several temperature controllers—top, bottom and side-draws. Occasionally, side-draws are found in other processes on true distillation columns, as well. Another special case, although it’s also a trick answer, is dualtemperature control where the would-be-ratio controller is replaced with a proportional-only temperature controller with feedforward signal. This provides for minor adjustment based on temperature deviation, but without the integral action that leads to wind-up and instability (although it is still easy to get instability if gain is set too large). The feedforward signal serves the role of the original ratio control. In my experience, this does not improve matters noticeably, but it can please other stakeholders (who don’t “get” the limitations) to see the controller at least make an effort, and it makes it easy to switch between top and bottom ratio control, especially without impacting data historians, documentation and graphics, since the temperature control tags persist either way. A final point in the debate has been model-based multivariable control (MPC) and whether it changes the game. Arguably, where dual-temperature control is feasible, model-based control has a better shot of keeping it in the window, based on its awareness of all the variables affecting separation (not just the relative gains of the two temperature controllers). But my experience is that neither MPC nor inferentials (which are essentially temperatures) change the (un)likelihood of success and many late-model MPCs exhibit similar modern day wreckage in the form of the reflux or reboiler MVs that are “clamped” or “out-of-service.” Some things never change! HP 1

2 3 4

LITERATURE CITED Friedman, Y. Z., “Distillation column dual-temperature control, Hydrocarbon Processing, March–April, 2010, (letters, August 2010). Kern, A. G., “More on APC designs for minimum maintenance,” Hydrocarbon Processing, December 2009. Friedman, Y. Z., “APC designs for minimum maintenance,”Hydrocarbon Processing, June–August, 2009 (and letter by P. R. Latour, October 2009). Greg, F., “Multivariable control of distillation,” ControlGlobal.com, May–July, 2009.

Tim is HP’s European Editor and has been active as a reporter The Lloyd authorWright has 30 years of process control experience and has authored sevand conference chair in the European downstream industry since 1997, before eral papers on inferentials, expert systems, advanced controls and decision support which he was a feature writer and reporter for the UK broadsheet press and BBC systems, with emphasis on practical process control effectiveness. Mr. Kern is a radio. Mr. Wright lives in Sweden and is founder of a local climate and sustainability professional engineer (inactive), a senior member of ISA, and a 1981 graduate of initiative. the University of Wyoming. HYDROCARBON PROCESSING NOVEMBER 2010

I 15


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LCFS will adversely affect oil-sands crude refiners Purvin & Gertz, Inc., has released a report on the effect low-carbon-fuel standards (LCFS) will have on oil sands. The report notes that LCFS programs are being implemented in California, Oregon and British Columbia. They are under consideration in many other states and provinces and are becoming regional in nature. LCFS programs differ by jurisdiction, but have in common mandated reductions in the carbon intensity of transportation fuels. By targeting petroleum-derived gasoline and diesel and promoting low-carbon alternative energy forms such as electricity, hydrogen, natural gas and next generation biofuels, LCFS programs are intended to reduce overall greenhouse gas (GHG) emissions on a “well-to-wheels” basis. “In effect, LCFS programs contribute to an ‘off-oil’ strategy,” said Tom Wise, who directed the study. The energy needed to produce oil sands crudes is higher than for most conventional crudes, so the resulting carbon intensities of refinery-produced gasoline and diesel from oil sands are also higher. Tom Wise points out it is a mistake to paint all oil sands crudes with the same brush because there are different oil sands crudes, such as synthetic crude oil and bitumen blended with various diluents, and each has a different pathway and carbon intensity. “Contrary to widely held perceptions, our study concludes that some oil sands diluted bitumen does not have high carbon intensity under the California regulations and should not carry an LCFS penalty,” Mr. Wise said. The Purvin & Gertz study estimates the well-to-wheels carbon intensities of refinery-produced gasoline and diesel from various oil sands and conventional crude oils. The study estimates the impact on consumer product prices, refinery margins and oil sands crude prices, for a range of LCFS carbon costs. In market regions that implement LCFS programs, consumer product prices will increase and refinery margins will fall.

“Some of the oil-sands crudes would require price discounts to compete with conventional crudes due to a reduced incentive to refine or upgrade heavy crudes,” said Mr. Wise. Further, reduced crude runs in market regions with LCFS programs could result in refinery closures and displace oil sands crudes to other markets. For instance, LCFS programs in the US Midwest would cause leakage of oil sands crudes to the US Gulf Coast or Asia-Pacific.

US demand for lubricants to reach 2.25 billion gallons in 2014 US demand for lubricants is forecast to expand 1.3% annually to 2.25 billion gallons in 2014, valued at $22 billion. This represents a significant improvement over the performance of the 2004–2009 period, when lubricant demand declined 5% annually. A turnaround in motor vehicle production, along with an acceleration in the number of automobiles in use, will support demand for automotive lubricants. Additionally, increased manufacturing output will drive demand for industrial lubricants. However, total lubricant consumption is not expected to reach pre-recession levels. This will largely be due to the greater use of longer-lasting, higher-performing synthetic lubricants that extend drain intervals, therefore reducing overall lubricant requirements in volume terms. Average price increases will continue to be significant due to expected growth in crude oil prices and a shift in product mix toward higher-value lubricants. These and other trends are presented in a new study from The Freedonia Group, Inc. Engine oils accounted for more than half of total US lubricant demand in volume terms during 2009. A significant rebound in motor vehicle output following the double-digit annual declines of the 2004–2009 period will propel engine oil demand in the factory-fill segment. However, this represents only a small fraction of engine oil demand, and the overall outlook for these products will be restricted by lengthening oil change intervals and the use of high

performance synthetic lubricants. As such, aftermarket demand will decline, with the “do-it-yourself” segment continuing to lose out to “do-it-for-me” services, a trend that stalled in 2008 and 2009 as drivers sought out more economical alternatives for their vehicle service needs during difficult economic times. Process oils—including white oils, rubber oils, electrical oils, ink oils, agricultural spray oils and defoamer oils—represent another leading lubricant category. Demand for these products is forecast to advance at the most rapid pace, promoted by rebounding manufacturing activity following the real (inflation-adjusted) declines of the 2004–2009 period. In particular, an improved outlook for food and beverages, chemicals, plastics and rubber will offer good growth opportunities. However, process oils will continue to encounter challenges brought about by changing environmental and regulatory standards.

Linde supplying hydrogen technology to US BMW plant The Linde Group will provide the BMW Manufacturing Co. plant in Spartanburg, South Carolina, with a hydrogen fueling system for its material-handling fleet. An according agreement was signed in mid-August 2010. More than 85 pieces of material-handling equipment are having their lead acid batteries replaced with hydrogen fuel cells. The trucks deliver process parts to assembly areas throughout the plant. After the conversion, this part of BMW’s internal logistics will be completely emission-free. “This is one of the largest hydrogen applications of its kind,” said Dr. Andreas Opfermann, head of innovation management of The Linde Group. “We are proud to work together with BMW, supplying both a high-efficiency fueling system and hydrogen with almost no carbon footprint. This project clearly shows the potential that hydrogen offers for internal logistics.” Refueling at the six indoor dispenser stations will be facilitated by Linde’s ionic compressor fueling system, combining efficiency and high throughput with lowHYDROCARBON PROCESSING NOVEMBER 2010

I 17


HPIMPACT maintenance costs and low noise compression. The hydrogen supplied to BMW, a byproduct from a chemical plant, is purified, compressed and liquefied by Linde using electricity produced from renewable hydropower. “We have a clear vision and we are determined to reach our goal of using renewable energy as much as possible throughout the plant site,” said Josef Kerscher, president of BMW Manufacturing. “We realize this is

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Oil and gas reserves increased 3 percent in 2009 The worldwide upstream investment of 224 oil and gas companies decreased 23% to $378 billion in 2009, according to a report released by IHS Herold. Although development spending fell nearly 20%, the first decline in a decade, total hydrocarbon reserves increased 3% as both oil and gas reserves grew for the first time since 2005. Production also increased 1%, driven by a 2.2% increase in natural gas output. “We were very surprised at the strength of reserve additions given the weak economic conditions and tightness in credit markets during 2009,” said Nicholas Cacchione, director of IHS Herold and author of the report. “As an industry, we spent fewer dollars, but they went further in terms of purchasing power.” Oil reserves reversed a two-year decline, rising 3% to 164 billion barrels. The main driver was 8.6 billion barrels in positive reserve additions, but extensions and discoveries in the Canadian oil sands and South and Central America also added a record 7.9 billion barrels. Natural gas reserves climbed 3.7% despite a record 11.4 trillion cf in negative reserve revisions, as development of unconventional plays in North America and liquefied natural gas (LNG) resources in Asia accelerated. The decline in capital spending was led by a 40% reduction by E&P companies, while the integrated oil companies cut investment by just 9%. Exploration spending was most resilient, dropping just 12% to $62.7 billion. In contrast, unproved acquisition costs were down 71%, and a 2% dip in proved acquisition outlays would have fallen 50% were it not for the $20 billion Suncor/Petro-Canada merger. “With the recession and ongoing uncertainty in the market last year, companies put acreage acquisition on hold and seemed to focus on their in-house development opportunities,” said Mr. Cacchione. “This decision, I think, reflected their desires to monetize known holdings that can be brought into production much more rapidly than something with a less certain payout several years down the road.”


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HPIMPACT Lower capital spending and higher reserves resulted in a near 50% decrease in reserve replacement costs (to $11.41/boe) and lowered finding and development costs to $12.23/boe. Strong natural gas reserve additions led reserve replacement rates to the highest levels in five years. Despite the strong performance metrics, upstream profits plunged 47% as a 13% decline in pre-tax expenses did not offset a 30% reduction in revenues. The integrated

oil companies accounted for 85% of the universe’s profit with the E&P companies accounting for the balance. Reserve writedowns slashed net income for the large E&P companies and drove the mid-sized and small E&P companies to a loss. However, the industry generated free cash flow due to the steep decline in capital investment. The IHS Herold report found that dividends rose modestly to another record level, which it noted “is remarkable” given

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the turmoil in the financial markets and the generally miserable results in the industry’s downstream operations. Dividends exceeded $100 billion, but common share repurchases were 23% lower, falling for the first time since 2004. Capital constraints brought about by reduced revenue and rising costs have almost completely eliminated share buybacks as a viable use of funds. Key regional findings of the report include: • Strong drill-bit additions aided improving results for reserve replacement costs and rates in the US. Unit profitability declined for the fourth consecutive year. Mineable bitumen reserve additions in Canada offset weak natural gas reserve additions. Profits were down sharply in Canada as well. • Oil and gas reserves in Europe continued to decline as companies redirect cash flows to other regions. The reserve replacement rate reached a five-year high through improving reserve additions, but the region was still below full reserve replacement figures. • Capital spending in the Africa and Middle East regions was down 14%, which was much less than the worldwide average. This drop in spending is due to regional dominance by the integrated oil companies, that tend to spend through the commodity cycles. • Asia-Pacific reserves gained 3% as natural gas extensions and discoveries surged. Reserve replacement rates in the region were well above full replacement levels. • Capital spending in South and Central America increased since regional players have strong development portfolios to exploit. Total reserves in the region increased 3%, the first gain in several years. • An uptick in proved acquisition spending limited total capital spending decline to 17% in the Russia/Caspian region, while drill-bit spending outlays fell 22%. Production in the region increased 18%, with strong results from both oil and gas output. IHS Herold anticipates a modest global rebound in upstream capital spending in 2010. “In North America, E&P investment increased 30% in the first half of 2010, which was higher than expected,” Mr. Cacchione said. “We think this should drive a global investment increase of more than 20% for the year. Outside of North America, where spending declines were less severe, we foresee upstream investment rising about 10%.” HP


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New FCC catalyst announced Grace Davison has released a new fluid catalytic cracking catalyst technology called Alcyon. This catalyst reflects Grace’s continuing commitment to delivering a broad portfolio of FCC catalysts and additives. Alcyon is designed for applications that require maximum activity and controlling coke make. Alcyon catalyst contains a proprietary zeolite modification that delivers the highest activity of any catalytic technology with a low surface area per kinetic conversion. An added benefit of the new zeolite is superior activity retention that has been verified in field performance. At constant surface area, Alcyon is more active than a traditional catalyst; thus it requires less catalyst-to-oil (C/O) ratios to achieve desired conversion levels and products. This new FCC catalyst uses improved access to active sites with increased cracking rates per active site. It is the latest of 11 new refining products introduced since 2008 by our research and development group,” commented Shawn Abrams, vice president and general manager of Grace Davison Refining Technologies. “Like all our FCC catalysts and additives, Alcyon catalyst’s revolutionary technology and performance are supported by our strong technical service and flexible manufacturing system.” Alcyon is particularly well suited for refiners looking to reoptimize unit operation to maximize profitability, yet remain within operating limits. For any given coke yield, Alcyon catalyst achieves higher conversion, maximizing total FCC barrels to the refinery gasoline pool. Select 1 at www.HydrocarbonProcessing.com/RS

Collaboration improves emergency valve performance ABB, the leading power and automation technology group, announced that it will collaborate with Dresser Masoneilan, a global leader in process control valves, on an integrated process to monitor, test and manage emergency shutdown valves (ESDVs) during all operational conditions, from normal plant operations to abnormal situations. These valves are crucial process elements for the oil, gas and

petrochemical industries, as well as for many other industrial processes. “Our collaboration with industry leaders like Dresser Masoneilan helps us to offer our mutual customers best-in-class safety solutions that will protect the integrity of their processes and the surrounding community,” said Luis Duran, Americas business development manager for Safety Systems, ABB. The combined solution leverages the capabilities of ABB’s 800xA high-integrity safety instrumented system (SIS) and Masoneilan’s SVI II emergency shutdown device (ESD) and PST controller to improve overall plant safety and increase the availability of ESDVs for optimal response of the isolation valve in emergency situations. This integration also simplifies safety compliance by automatically recording partial stroke test results and emergency shutdown events, saving time and money while increasing efficiency. “By taking advantage of System 800xA’s unique integration capabilities and open standards, the user has immediate access to the health diagnostics and status of the emergency shutdown valve. This access also provides proactive management of this critical device, for instance, enabling remote triggering of partial stroke tests, to ensure that it is ready to perform when needed,” said Kristian Olsson, manager of ABB’s Safety Center of Excellence. “This immediate readiness is vital to the protection of the process, the environment and the surrounding community in the event of an abnormal situation.” As an integrated object within System 800xA, Masoneilan’s SVI II ESD device can be configured to perform scheduled partial valve stroke tests while remotely monitoring and maintaining the emergency shutdown valves during normal plant operations. This minimizes the need for outages and downtime to evaluate the health and readiness of these critical process elements. This also provides easy-tounderstand alerts and recommendations regarding valve status, as well as required partial stroke test and emergency shutdown signatures and documentation. “While open standards offer great benefits for end users, it is the collaboration

between automation vendors that provides for an ‘out-of-the-box’ solution capable of generating instant results,” said Sandro Esposito, global marketing manager of digital products for Dresser Masoneilan. “The SVI II ESD provides an excellent return on investment with its combined shutdown function, partial stroke test function and shutdown event ‘blackbox’ into a single SIL3-certified device.” The SVI II ESD is the latest technology in emergency shutdown valve automation and in-service valve partial stroke testing. The SVI II ESD is the only SIL3-certified ESD certified at 4mA with stainless steel housing. The device can be implemented using a 4/20mA signal (analog safety demand), 0-24Vdc (discrete safety demand) or a combination of both. Standard on the device are an LCD display and explosion-proof external pushbuttons. This design architecture offers a sophisticated platform while being Type A (simplex device) compliant. System 800xA high integrity is ABB’s next generation safety system. This SIL 3-rated SIS provides the highest level of integration of safety and control on the market and a unique embedded diverse technology architecture that provides superior protection of the process, plant, personnel and the environment while it optimizes overall process efficiency. Select 2 at www.HydrocarbonProcessing.com/RS

Gasoline benzene reduction alkylation catalytic technology Today, more stringent regulation is challenging refiners to meet the tightening gasoline specifications for benzene at As HP editors, we hear about new products, patents, software, processes, services, etc., that are true industry innovations—a cut above the typical product offerings. This section enables us to highlight these significant developments. For more information from these companies, please go to our website at www. HydrocarbonProcessing.com/rs and select the reader service number.

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HPINNOVATIONS olefin streams to remove potential contaminants such as sulfur and nitrogen species. Some process advantages include: Benzene reduction—Reduction of benzene content in the gasoline pool to meet new benzene regulation. High benzene conversion can be achieved. Gasoline volume swell—Upgrading of light olefins and benzene into high-octane gasoline blend stock also results in a volume swell of the gasoline volume. The

specific volume swell will be dependent on the feed composition and level of benzene conversion. Octane gain—2–5 numbers of (R+M)/2 increase is typical. The specific octane gain depends on the feed composition. Reformer flexibility for increased hydrogen production. Select 3 at www.HydrocarbonProcessing.com/RS

Simulation to Business. Knowledge to Profit.

Do all safety valves open simultaneously? Dynamic Simulation following the API guideline can help you to reduce Capex during flare system debottlenecking accounting for the real transient behaviour of your plant during emergencies or scheduled blow-downs. SIMULTANEITY EFFECTS OF A NUMBER OF PROCESS UNITS

Sum (t) FCC Flare load, t/h

the lowest cost and without significant octane loss. For instance, the US Environmental Protection Agency’s most recent clean fuels regulations (Mobile Source Air Toxics II) require refineries to reduce benzene to less than 0.62 vol% (on an annual basis) in gasoline by 2011, from its current level of 1.0 vol%. The limits apply to both reformulated and conventional gasoline. Sources of benzene in the gasoline pool vary for each refinery; the predominant source for most is reformate. Benzene reduction in reformate can be achieved by either the removal of benzene precursors before they are converted to benzene in the reformer, or by post-removal of benzene in the reformate product by chemical conversion of benzene or by removal with fractionation. EMRE developed BenzOUT, a reformate alkylation technology, to convert benzene into high-octane alkylaromatic compounds (such as isopropylbenzene) for gasoline blending by reacting a benzene-rich stream with light olefins such as propylene in low-value olefin streams. This patented technology avoids the octane loss and hydrogen consumption associated with the alternative option of benzene saturation. BenzOUT is a refining process based on ExxonMobil’s ethylbenzene and cumene technologies, which have been widely used in the chemical industry with over 80 worldwide commercial applications. In addition, EMRE’s BenzOUT technology was commercially demonstrated at a North American refinery. BenzOUT reduces benzene by reacting a benzene concentrate stream with a light olefincontaining stream such as LPG over a proprietary catalyst. Typically, a benzene concentrate ranging from 10% to 50% is processed. Key features of the process include: Fixed-bed catalyst technology—This process uses a simple fixed bed reactor. In revamp projects it is possible to retrofit existing polygas tubular/chamber reactors or spare reformer reactors for this application. Catalyst—This process utilizes a proprietary solid acid catalyst with long cycle lengths. It possesses high catalyst activity and allows long catalyst cycle life. In addition, the catalyst is completely regenerable ex-situ to further extend catalyst life. Feed requirement—This process requires conventional feed pretreatment for the

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

North America Flint Energy Services Ltd. was awarded a contract extension for work on Suncor Energy’s Firebag SAGD projects near Fort McMurray, Alberta, Canada, for site-wide construction and commissioning services. The contract extension of approximately $78 million will increase the company’s total contract scope to just over $110 million. Work on the contract will be completed by Flint’s Facility Infrastructure Division and will employ over 700 workers through early 2011. KP Engineering, LP, recently completed a 13,000-bpd FCC gasoline hydrotreater for Wynnewood Refining Co. in Wynnewood, Oklahoma. The project, which was executed through an engineering, procurement and construction management (EPCM) contract, achieved mechanical completion in September. The hydrodesulfurization unit utilizes the Prime-G+ process, licensed from Axens North America, Inc. The unit is designed to produce FCC gasoline with a sulfur content of less than 50 ppm and is a key component of Wynnewood’s overall strategy for clean transportation fuels. Unit startup is scheduled for the fourth quarter of 2010. Gulf Coast Fractionators, a partnership among ConocoPhillips, Devon Energy Corp. and Targa Resources Partners LP, announced plans to expand the capacity of its natural gas liquids fractionation facility located in Mont Belvieu, Texas. The maximum gross fractionation capacity of the facility will be expanded by approximately 42% (43,000 bpd) to 145,000 bpd. ConocoPhillips, as the operator, will manage the expansion project and existing operations are not expected to be disrupted during the construction phase. The expansion is expected to be operational during the second quarter of 2012, subject to regulatory approvals. The total capital expenditures of approximately $75 million are expected to be significantly lower than a greenfield fractionation facility since the new capacity will be integrated with the existing fractionation capacity, utilities, infrastructure and footprint already at Mont Belvieu.

BlueFire Renewables, Inc., has finalized and signed an engineering, procurement and construction (EPC) contract for its planned cellulosic ethanol facility in Fulton, Mississippi. The facility will be engineered and built by Wanzek Construction, Inc., a wholly owned subsidiary of MasTec, Inc., for a fixed price of $296 million which includes an approximately $100 million biomass power plant as part of the facility. The contract is negotiated in a manner to be appealing for non-recourse project bank financing and, more importantly, serves as the final key project contract agreement to move forward with both the DOE and USDA loan guarantee programs. The project will allow BlueFire to utilize green and wood wastes available in the region as feedstock for the ethanol plant that is designed to produce approximately 19-million gpy of ethanol. Technip has two lump-sum contracts with Valero Refining Co. and Diamond Shamrock Refining Co. for two hydrogen plants at the company’s refineries in Memphis, Tennessee, and McKee, Texas. The TREND ANALYSIS FORECASTING two 30 million standard cubic feet per day Hydrocarbon Processing maintains an hydrogen plants will produce high purity extensive database of historical HPI projhydrogen and export steam. Theactivity plants ect information. Current project is published three times a top-fired year in thesteam HPI will use a high efficiency Construction Boxscore. When project methane reforming process and autilize the is completed, it is removed from current latest nitrogen oxide reduction technology listings and retained in a database. The thereby ensuring minimum emissions. database is a 35-year compilation of projects type, operating company, licenThebycontracts cover basic engineering, sor, engineering/constructor, etc. project management, detaillocation, engineering, Many companies use the historical data for fabrication, supply and installation, pretrending or sales forecasting. commissioning and startup assistance. The historical information is available in Technip’s operating center in Claremont, ® and comma-delimited or Excel can be cusCalifornia, execute theseThe contracts. tom sorted will to suit your needs. cost of the sort is depends on the size Performance and complexTechnip partnered with ity of the sort you request and whether a Contractors Inc. for the installation customized program must be written. You ofcan both hydrogen plants. Thesuch project focus on a narrow request as theis scheduled completed first half history oftoa be particular typein ofthe project or can obtain the entire 35-year Boxscore ofyou 2012. database, or portions thereof.

Simply send Group, a clear description of the data Willbros Inc., has been selected you need and you will receive a prompt bycost Diamond Shamrock quotation. Contact: Refining Co. to

provide engineering, procurement and conLee Nichols struction (EPC) of the P. O. Box new 2608flare gas recovHouston, 77252-2608 ery facilities at theTexas, Valero McKee refinery Fax: 713-525-4626 located in Sunray, Texas. The project is valLee.Nichols@gulfpub.com. ued ate-mail: approximately $14 million.

South America GE Oil and Gas has been awarded a $32 million contract by PlusPetrol Peru Corporation SA to supply compression and reinjection upgrade equipment to boost the output of the Malvinas natural gas liquid (NGL) processing plant, located 400 km from Lima at the southern section of the Peruvian Amazon rainforest. Under the contract, GE will upgrade two existing GE gas turbine modules with GE technology to reduce emissions and other technology to reduce outages and extend the life cycle of the equipment. In addition, three existing modules will be upgraded in a different fashion. The equipment was manufactured in Florence, Italy, and shipped to Peru in March 2010. With the use of trucks prohibited in the environmentally sensitive region, GE overcame the logistical challenge of “on-time” delivery by taking advantage of seasonal swelling of the Amazon and Urubamba rivers to transport the two 60-ton modules to the Malvinas plant. Foster Wheeler AG’s Global Engineering and Construction Group has been awarded a contract by YPF SA for a new delayed coking unit to be built at YPF’s Complejo Industrial La Plata in Argentina. Foster Wheeler’s scope of work includes TREND ANALYSIS FORECASTING Hydrocarbon Processing maintains an extensive database of historical HPI project information. The Boxscore Database is a 35-year compilation of projects by type, operating company, licensor, engineering/constructor, location, etc. Many companies use the historical data for trending or sales forecasting. The historical information is available in comma-delimited or Excel® and can be custom sorted to suit your needs. The cost of the sort depends on the size and complexity of the sort you request and whether a customized program must be written. You can focus on a narrow request such as the history of a particular type of project or you can obtain the entire 35-year Boxscore database, or portions thereof. Simply send a clear description of the data you need and you will receive a prompt cost quotation. Contact: Lee Nichols P.O. Box 2608, Houston, Texas, 77252-2608 Fax: 713-525-4626 e-mail: Lee.Nichols@gulfpub.com HYDROCARBON PROCESSING NOVEMBER 2010

I 27


HPIN CONSTRUCTION detailed engineering, procurement services and support to construction and plant startup. The unit will use Foster Wheeler’s delayed coking technology. The planned new facility will have a capacity of 28,000 bpsd, and will produce anode coke to be used in the aluminum industry.

Europe Avantium has started with the construction of a pilot plant at the Chemelot site in

Geleen, The Netherlands. The pilot plant will produce building blocks for making green materials and fuels. Avantium has developed a proprietary catalytic process to convert carbohydrates into furanic building blocks under the brand name YXY. The pilot plant is expected to become operational in the first quarter of 2011. It will demonstrate the process developed in Avantium’s labs on a larger scale. Furthermore, it will produce several tons of

furanic building blocks per year to support product development. Avantium is collaborating with industrial partners such as NatureWorks (a subsidiary of Cargill) and Teijin Aramid to develop novel materials on the basis of its furanic building blocks. Avantium collaborates with DAF Trucks (a Paccar company) on the development of furanic building block fuels. The pilot plant is partly funded by a €1 million grant from the Dutch Ministry of Agriculture.

Middle East Siirtec Nigi S.p.A. has been awarded two contracts by Saipem S.p.A. to design and supply eight packages, each including a burner, reaction furnace, steam drum, waste heat boiler, and 20 sulfur condensers. This equipment represents the core of the sulfur Claus unit. It will be installed at the Shah gas field in the UAE. The field units will process around 1 billion scfd of sour gas to produce around 540 million scfd of gas suitable for consumption. Siirtec Nigi’s scope of work includes the supply of know-how, design and supply of the equipment together with the relevant process guarantees. In addition, Siirtec Nigi will provide site services and specialists for commissioning and startup.

Asia-Pacific

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Select 155 at www.HydrocarbonProcessing.com/RS

Toyo Engineering Corp. has a contract for an ethanolamine (ETA) production plant to be constructed in Jiaxing, Zhejiang Province, China, by Honam Petrochemical Corp. (HPC). The plant is scheduled for completion in the first quarter of 2012. The ETA project is the third contract that Toyo has received from HPC, following a polypropylene plant project awarded in 1977 and an ethylene plant project in 1989. The past successful projects and supports from Toyo Group, including ToyoKorea, have led to the award. The plant will produce ETA. As a result of bidding, Toyo-China has won the engineering, procurement and construction (EPC) contract after successfully competing against other contractors in China and Korea. China Petroleum and Chemical Corp. plans to build a liquefied natural gas (LNG) receiving terminal in Maoming, Guangdong Province, China. The construction project will include an oil and gas wharf, an LNG receiving terminal and 500,000-cubicmeter LNG storage tanks. HP


HPI CONSTRUCTION BOXSCORE UPDATE Company

City

Plant Site

Project

Capacity Unit Cost Status Yr Cmpl Licensor

Sonatrach Kenya Petroleum Refineries NNPC

Algeries Mombasa Kogi

Algeries Mombasa Kogi

Refinery Refinery Refinery

RE RE TO

60 Bcfd None 750 bpd

908 M 2013 H 23000 E

Queensland Anning Anning Zhanjiang Bhatinda Panipat Paradeep Visakhapatnam

Bowen-Surat Basin Kunming Oil Refinery Kunming Oil Refinery Zhanjiang Bhatinda Panipat Paradeep Visakh Refinery

LNG Offsites Refinery Refinery Refinery Hydrocracker LPG Terminal Refinery

EX TO TO EX TO EX EX TO

1.5 200 200 350 9 1.7 15 300

750 3400 3400 7600 189 950 900 10000

Ocean Cay Ocean Cay Ocean Cay Pernambuco Barrancabermeja Barrancabermeja Barrancabermeja Cartagena Kingston Minatitlan Jose Anzoategui

Ocean Cay Ocean Cay Ocean Cay Pernambuco Barrancabermeja Barrancabermeja Barrancabermeja Cartagena Kingston Minatitlan Jose Anzoategui

Desalination LPG Terminal Storage, LPG Refinery Hydrogen Sour Water Stripper Treater, Tail Gas Refinery Gas Plant Refinery Methanol

Ras Laffan Ras Az Zawr Yanbu Yanbu Yarimca Ruwais Ruwais

Ras Laffan Ras Az Zawr Yanbu Yanbu Yarimca Ruwais Ruwais

Gas to Liquid Desalination Hydrogen Silicones Refinery Refinery Refinery (2)

bpd MIGD Mcfd t/a m-tpy bbl bpsd

24000 1460 450 1100 10000 623 3110

Lemont Wabash Beaumont Sunray

Lemont Wabash Beaumont Sunray

Hydrotreater, ULSD 42500 bpd Coal Liquefaction (CTL)TO 2.5 m-tpy Isomerization EX 600 gpm Steam Methane Reformer EX 30

451 3000 1 Scfd

Engineering Constructor

AFRICA Algeria Kenya Nigeria

Technip

ASIA/PACIFIC Australia China China China India India India India

Santos\PETRONAS JV CNPC CNPC Sinopec HMEL Indian Oil Corp Ltd IOCL HPCL

Mtpy Mbpd Mbpd bpd m-tpy MMtpy m-tpy bpd

P U U U U U P U

2011 2012 2012 2014 2011 2010 2012 2014

Total

UOP

EIL

EIL

Technip KTI Tipiel|Axens Prosernat KBR|KBC|EMRE|UOP|Lummus Technology UOP

Axens Tipiel|Axens Axens Technip|CB&I

Sadeven|Ismocol Sadeven|Ismocol Sadeven|Ismocol CB&I

LATIN AMERICA Bahamas Bahamas Bahamas Brazil Colombia Colombia Colombia Colombia Jamaica Mexico Venezuela

MIDDLE Qatar Saudi Arabia Saudi Arabia Saudi Arabia Turkey UAE UAE

UNITED Illinois Indiana Texas Texas

AES Corp AES Corp AES Corp Petr Brasileiro SA Ecopetrol Ecopetrol Ecopetrol Ecopetrol Petrojam Ltd Pemex PDVSA

ST

TO

EX TO RE

500 Mgpd P 2012 None P 2012 40 Mm3 P 2012 230 bpd 12000 U 2011 19 MMscfd 30 C 2010 500 gpm 21 C 2010 110 t/a 17.4 C 2010 165 Mbpd 2600 P 2013 7 Mbpd 1300 F 2013 240 bpd 318 P 2011 850 m-tpy 7 C 2010

Tecnicas Reunidas Constr N. Odebrecht

EA

Shell Royal Dutch Doosan Saudi Aramco Al Rajhi Petrochemical Petkim Petrokimya Hldg Takreer Takreer

ATES

TO TO

140 228 262 6 TO 10 EX 1271 EX 127200

U U U P U P E

2012 2014 2014 2014 2015 2014 2014

Descon Eng

JGC|KBR|McDermott Air Liquide

GS E&C Bayer|Lummus Technology|Axens|Shaw S&W GS E&C

GS E&C GS E&C

ST

Citgo Clean Coal Refining Corporation Arabian American Development Valero Refining Co

C 2010 S 2011 C 2010 14 U

Mustang 2012

Turner Technip

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

BOXSCORE DATABASE

ONLINE

THE GLOBAL SOURCE FOR TRACKING HPI CONSTRUCTION ACTIVITY For more than 50 years, Hydrocarbon Processing magazine remains the only source that collects and maintains data specifically for the HPI community, publishing up-to-the-minute construction projects from around the globe with our online product, Boxscore Database. Updated weekly, our database helps engineers, contractors and marketing personnel identify active HPI construction projects around the world to: • Generate leads • Market research • Track trend analysis • And, decide future budget planning. Now, we’ve made our best product even better! Enhancements include: • Exporting your search results to Excel so you can compile your research • Delivering the latest updated projects directly to your inbox each week • Designing customized construction reports for your company using our 50 years of archived projects. For a Free 2 -Week Trial, contact Lee Nichols at +1 (713) 525-4626, Lee.Nichols@GulfPub.com, or visit www.ConstructionBoxscore.com

Select 156 at www.HydrocarbonProcessing.com/RS HYDROCARBON PROCESSING NOVEMEBR 2010

I 29


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HPI VIEWPOINT Stretch in technology and gaps in process safety for the hydrocarbon industry Dr. M. Sam Mannan, PE, CSP is a Regents Professor in the Chemical Engineering Department at Texas A&M University and director of the Mary Kay O’Connor Process Safety Center at the Texas Engineering Experiment Station. The mission of the Center is to improve safety in the chemical process industry by conducting programs and research activities that promote safety as second nature for all plant personnel in their day-to-day activities. Before joining Texas A&M University, Dr. Mannan was vice president at RMT, Inc., a nationwide engineering services company. Dr. Mannan is a registered professional engineer in the states of Texas and Louisiana and is a Certified Safety Professional. His experience is wide ranging, covering process design of chemical plants and refineries, computer simulation of engineering problems, mathematical modeling, process safety, risk assessment, inherently safer design, critical infrastructure vulnerability assessment, aerosol modeling, and reactive and energetic materials assessments. Dr. Mannan co-authored the Guidelines for Safe Process Operations and Maintenance, published by the Center for Chemical Process Safety, American Institute of Chemical Engineers (AIChE). He is the editor of the 3rd edition, Lees’ Loss Prevention in the Process Industries; he has published 137 peer-reviewed journal publications, two books, seven book chapters, 142 proceedings papers, 12 major reports, and 143 technical meeting presentations. Dr. Mannan is the recipient of numerous awards and recognitions including the American Institute of Chemical Engineers Service to Society Award, the Texas A&M University Association of Former Students’ Distinguished Achievement Award for Teaching, the Texas Engineering Experiment Station Research Fellow, the Texas A&M University Dwight Look College of Engineering George Armistead, Jr. ’23 Fellow. In 2007, he was elected Fellow of the AIChE. In December 2008, the Board of Regents of Texas A&M University System recognized Dr. Mannan’s contributions in teaching, research and service by naming him Regents Professor of Chemical Engineering. Dr. Mannan received his BS degree in chemical engineering from the Engineering University in Dhaka, Bangladesh. He obtained his MS and PhD in chemical engineering from the University of Oklahoma.

Oil and gas are contributing enormously to the quality of our lives in the 21st century, just as they were throughout the 20th century. With the economies in the various parts of the world expanding significantly, more energy sources are required by our society, and in the previous several decades, people have switched to offshore deepwater hydrocarbon reservoirs. Besides politics, current exploration and production are limited by the technology, for deeper wells, higher pressure reservoirs or crudes that are difficult to recover because of higher viscosity. Innovative offshore technology must be developed to carry out deepwater production and operations. At the same time, these hazardous operations (i.e., deeper wells and higher-pressure reservoirs) are creating new and unique hazards. Along with the current attention on the Transocean Deepwater Horizon Oil Spill, the development of advanced technologies to ensure the process safety and operational reliability of offshore facilities is becoming extremely important.

Hazards offshore. At present, we face the challenge of how to

prevent or control hazards of deepwater exploration and production. Process safety is always an essential part of the oil and gas industry and a core value that requires continual improvement. Regulations should be dynamic and ready to be modified based on occurring industrial issues. Accordingly, the government and offshore operators should develop comprehensive management programs or regulations that assess process safety and environmental hazards. It is well known that offshore operations have a very special environment, involving drilling, production and transport, as well as emergency response to incidents. Offshore employees are faced with many different factors that increase their exposure to injury, such as poor weather conditions, high-pressure operations, chemicals and confined space. Due to the dangerous nature of offshore operations, employees typically have very demanding work schedules. Research and technology developments for operation in deep waters and high-pressure reservoirs are urgently needed. The industry must develop theories, analytical techniques and technologies to improve the current offshore infrastructures from all sources of failure, including design, operations, management, natural disasters and intentional acts such as terrorism. The research should focus on theories and techniques that apply to the many process safety issues, such as structural integrity, layers of protection, offgas handling, drilling, risk assessment and consequence analysis, human error and safety culture. Test beds may include processing facilities and complex structures within the offshore infrastructure, transportation vehicles (e.g., ship, helicopter) and the marine environment. The goals of offshore safety research include integration of the concepts of process safety into the design and operation of offshore platforms and to use this to improve safety performance such that the unit/process is not vulnerable to certain failures. Also, by identifying aspects of the system that are vulnerable (not resilient), the speed and efficiency of response to failures could be improved. Some relevant areas of research include: Structural integrity of risers/pipelines. Risers and pipelines are subjected to different types of corrosion due to continuous loadings (fatigues) and their exposure to the marine environment. Research is needed to study corrosion behavior underneath the coating layer, to assess the integrity of particular structures and to develop coating/ corrosion assessment criteria for service under extreme conditions. Layers of protection. Many US offshore rigs are equipped with blowout preventer (BOP) casing shear rams to seal off an oil or natural gas well being drilled or worked on. However, the BOP is currently the only layer of protection within the system and it is vulnerable to single-point failure. Thus, research efforts are needed for developing/identifying multiple layers of protection. Risk assessment and consequence analysis using CFD. Computational fluid dynamics (CFD) has gained widespread recognition as a powerful tool for risk assessment and consequence analysis. The very nature of infrequent and highly diverse disasters can make prediction of the likelihood and consequences of such events very difficult. The multitude of systems and structures HYDROCARBON PROCESSING NOVEMBER 2010

I 31


HPI VIEWPOINT involved requires a very broad multidisciplinary team to understand, evaluate, compare and plan for disastrous failures. Research is needed with regard to fluid-structure interaction (vibrations of risers, motions of floating platforms), flow around vessel hulls in the presence of current and wind forces, wave loads (slam and impact), tank sloshing and BOP (impact) facility siting. Recovery of H2S. Hydrogen sulfide (H2S) is corrosive for carbon steels used in offshore structures and can lead to nervous disorders and acid rain. To remove H2S, a process called gas sweetening and amine sweetening has been used for onshore processing facilities. This process removes H2S from the feed and redirects it to other processes in the plant, where it is converted into elemental sulfur or sulfuric acid. However, the present process is limited by the operating conditions and requires additional units to be built, which makes it insufficient for the offshore facilities. Given the compact size of the offshore structures, there is a need to optimize the gas-recovery process for offshore facilities. Resilient operation of deepwater drilling. During drilling, all materials drilled out need to be removed, i.e., transported to the surface, a process that is referred to as hole cleaning. Often, some material remains in the well. Due to the number of parameters influencing hole cleaning and the complex mechanisms involved, the phenomenon has not yet been fully understood. Integration of the concept of resilience to the drilling and hole cleaning processes is needed. Human error and safety culture. Both human error and safety culture have been identified as contributing causes in indus-

trial accidents, including offshore facilities. Assessment of the safety culture in every aspect of work in the organization, and development of a systematic approach to apply safety culture in the organization will lead to reduction of human errors. Spills clean-up. Mechanical containment should be the pri-

mary line of defense against oil spills. However, this method is not effective in clearing a large spill area. Chemical and biological methods can be used in conjunction with mechanical means for containing and cleaning up oil spills. Dispersing agents are the most useful in helping to keep oil from reaching shorelines and other sensitive habitats. Biological agents have the potential to assist recovery in sensitive areas such as shorelines, marshes and wetlands. In this regard, research for new dispersants, including nano-surfactant technologies, is needed. Meeting the challenge. Deeper wells, higher pressure reservoirs and crudes with higher viscosity have introduced hazardous operations in offshore facilities, thus creating new and unique hazards. Efforts should be made to expanding the current focal point from that of drilling and utilization of BOPs to include a more encompassing proactive prevention program requirement for all offshore operations. There are many possible ways in which disasters equal to or greater than Deepwater Horizon could occur in offshore operations. Comprehensive programs that include inherent safety considerations, multiple layers of protection, consideration of human factors, analysis of worst-case scenarios and emergency response planning are needed. Finally, major efforts that explore new technologies in various areas should be a high priority. HP

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I NOVEMBER 2010 HYDROCARBON PROCESSING

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HPI VIEWPOINT Better water management: A crucial and growing requirement Glen Messina is global business leader, chemical and monitoring solutions —water and process technologies for GE Power and Water. In this role, he is responsible for setting the strategic direction of the chemical and monitoring solutions business as well as overseeing global execution. Mr. Messina leads the commercial and product management organizations as well as supporting functions. Previously, he was chief financial officer for water and process technologies and was responsible for leading the company’s finance function.

Water touches every industry and every sector, impacting lives, livelihoods, global industries and local economies. But the world is facing growing water scarcity challenges, both in developed and developing countries. Industry accounts for 20% of all water consumed globally. Refineries and petrochemical facilities, in particular, tend to be large water users, both for cooling purposes and in production operations. The hydrocarbon processing industry (HPI) can play a key role by adopting comprehensive water-management strategies and water reuse practices on a widespread basis. As the global economy slowly recovers, refineries can expect to see a shift in demand for finished petroleum products in mature markets and growing demand for fuels by developing regions. Globally, refining facilities will continually seek ways to increase production and profit margins, while reducing or eliminating unscheduled shutdowns and the subsequent untimely cost of cleaning and replacing equipment. Building a robust water-management strategy should help refineries to run safer, and more efficiently, and to be more profitable. Technology as an enabler. From a technology perspective,

water recovery, treatment, recycling and discharge solutions exist right now not only to enable conservation, but also to improve the productivity, efficiency, safety and overall economics of waterintensive industrial processes. Water technology companies offer technical services, knowledge and practical experience. For example, optimizing the performance of open recirculating cooling systems used in refineries and petrochemical plants represents an enormous opportunity to both lessen the impact on freshwater supplies, as well as to reduce costs, increase productivity and protect assets. These cooling systems rely on chemically treated water to reduce water consumption, to prevent deposits from compromising heat transfer in critical process equipment, to mitigate the effects of corrosion that threatens capital asset life, and to manage microbial growth, which, in addition to having a negative impact on production, can adversely affect human health. But the treatment and control process requires constant monitoring and finetuning, which until now has been cumbersome and inefficient. Advanced technology is altering that status quo. New technology gives users the ability to monitor, measure and adjust multiple chemicals in the cooling water in an online mode and more precisely than traditional methods. The result is a better-performing

cooling system with significantly lower operating costs, and at the same time, it saves hundreds of thousands of dollars in annual fresh water acquisition costs for a large system. Another area where advanced water-management technology plays a key role is to enable the increased use of wastewater in industrial processes. A case in point is a large refinery in southern Italy, which had to curtail its use of city water during a drought. If a second source couldn’t be found, the refinery would have had to shut down. A nearby municipal wastewater plant could provide the needed volumes, provided the water could be treated adequately. Mobile membrane-based reverse osmosis (RO) filtration technology was the solution. Once the municipal wastewater was run through the refinery’s own clarifier and filtration system, it was further treated by multiple RO filtration units until it was clean enough for refinery processes. Because these units were mobile and could be trucked to the site, they were able to be installed quickly. As drought conditions and demand continued to grow, the flexible nature of the system made it easy to add another unit without interrupting the production of treated water or refinery operations. Intimate technical understanding needed. A deep

technical knowledge of process chemistry is required to solve many of the difficult water-management issues faced by HPI facilities. For example, a Gulf Coast refinery was experiencing poor dehydration in the processing of heavy Venezuelan crude. Water carryover and salt content had increased to unacceptable levels. Although caustic had been injected to control overhead chlorides, desalter performance was impacting furnace-fuel usage, unit pressure drop and overhead corrosion control. A comprehensive analysis was done, and the ultimate solution was a new emulsion breaker that was injected into the crude charge. As time went on, the chemical injection rate was reduced to about 60% of the historical chemical injection dosages. Total chlorides were cut by 50% and more consistent corrosion control in the crude overheads was achieved, even with a reduction in caustic injection of nearly 30%. Tomorrow is upon us. Refining crude oil into high-value

fuels and other refined products requires careful coordination and precision of many complex and expensive systems. Profitability depends on maintaining the quality and efficiency of these key process units, controlling operating costs, meeting environmental requirements and improving efficiencies in every step of the production process, while ensuring final product quality. We have reached the point where the HPI can no longer assume the availability of adequate, uninterrupted freshwater supplies. Fluctuating water demand and supply increasingly will become the norm and the HPI must address growing environmental challenges and regulations. Partnering with suppliers who understand and can address concerns with sound water and process solutions will result in safer, more efficient, more profitable and more environmentally responsive operations. HP HYDROCARBON PROCESSING NOVEMBER 2010

I 33


Select 55 at www.HydrocarbonProcessing.com/RS


PLANT SAFETY AND ENVIRONMENT

SPECIALREPORT

Water among causes for storage tank explosion Reinvestigation uncovers true accident events M. FERJENCIK and B. JANOVSKY, University of Pardubice, Pardubice, Czech Republic

A

n explosion occurred inside a bitumen storage tank. An abrupt interruption of the purging steam discharge from the tank preceded the accident. The suspected ignition sources alone would not be able to cause the observed explosion. An additional impulse must have been present. Careful analysis of the event led to the conclusion that the explosion was probably preceded by water that was introduced inside the tank. Iron sulfides were present under the tank roof and combined with the injection of water (causing the fresh air suction and atmosphere movement inside the tank) and this explained the explosion under actual conditions. Introduction. Trevor Kletz wrote an

anecdote in which he exaggerated the dangerous properties of water.1 Nevertheless, he still described water as an extinguishing agent. However, in the accident that occurred, a situation encountered showed that water played the opposite role. Even introducing a relatively small amount of water into a large bitumen tank seems to have triggered the accident that resulted in an explosion and a fire. This article will present the accident’s reinvestigation. It is based on a report that was prepared by a plant investigation committee. However, the original report ignored the necessity to explain an abrupt interruption of purging steam discharge from the tank which preceded the accident. The reinvestigation concentrates on this event and makes it the focal point to shed light on why the accident occurred. Installation. The accident occurred more than five years ago in a plant that processed heavy-oil hydrocarbons. The event involved an almost 40-yr-old bitumen storage tank. The tank was an insulated, carbon-steel,

vertical cylinder with a fixed roof. It was standing in a concrete emergency sump. The tank communicated with the surrounding air via a couple of vents which were placed at the southern edge of its roof. The total tank volume was 1,200 m3—height 10 m, diameter 12.4 m (Fig. 1). The tank was equipped with openings and pipe connections. One hatch and seven pipe connections were in the bottom part of its wall. There were two connections of the steam heating system that kept the bitumen inside the tank hot and liquid. The second hatch and seven other openings were in the roof of the tank, and one blind nozzle was in the upper part of the tank wall. Fig. 2 only shows the roof openings that were considered to play a role in the accident. The steam inlet (50-mm diameter) was located close to two vent nozzles (200-mm diameter each). A steam pipe with a 25-mm diameter was inserted into the steam inlet. Its end piece inside the tank was used for steam purging. It was less than 1.5-m long; its mouth was blind and its wall was perforated. The bitumen circulation inlet pipe nozzle (125-mm diameter) position was across the roof at the northwest edge. The bitumen circulation inlet—a carbon steel pipe, 100-mm diameter was inserted—was the newest tank opening. It was welded during a scheduled outage of the tank, roughly six weeks before the accident. The second most recent opening was two years old.

ing was to be used only when the bitumen temperature inside the tank was higher than 190°C. More than one day before the accident, the bitumen level was 65 cm and at 155°C. These parameters were kept for more than 24 hr. Conditions inside the tank started to change again approximately 3 hrs 40 minutes before the accident. At that time, 120 ton batch of bitumen was transported into the storage tank. Since the temperature of the influent bitumen was higher than 200°C (up to 250°C), steam purging was commenced. The transport was finished 40 min before the accident— the bitumen level was 180 cm and at 205°C. Steam heating and steam purging were left in operation; steam pressure was about 4.5 bar and temperature about 140°C. The filling pipeline was emptied into the storage tank by pressurized air. All manipulations connected with the transport were finished 25 min before the accident. About 25 min later, an explosion occurred inside the bitumen storage, blowing off the tank roof. After the explosion, the tank contents started to burn. The roof fell beside the emergency sump and leaned against the southwest wall. The explosion did not damage the tank wall, but a subsequent fire destroyed the south-southwestern part. A layer of coke sediments were prob-

Chronological order. Operational

records showed that during the last four days before the accident, the bitumen level inside the storage tank was relatively low (maximum 330 cm) and the temperature was between 150°C–155°C. Steam heating was operated permanently and steam purging was inactive. It was in accordance with operational instructions: steam purg-

FIG. 1

Bitumen storage tank, front view.

HYDROCARBON PROCESSING NOVEMBER 2010

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SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

ably burned at relatively high temperatures. The fire was localized and extinguished in less than 5 hr. It did not cause material damage outside the bitumen storage tank nor any fatalities or injuries. Unburned bitumen had to be placed in barrels and reprocessed. No anomalies were found in the unburned liquid during its removal and reprocessing. Additional observations. The south-

ern part of the tank roof, with vent nozzles, was within the scope of one camera belonging to the plant security system. Relevant record analysis provided interesting additional information about what happened a few minutes before the accident. Records confirmed that emptying the filling line was performed within a 15-min time period, ending 25 min before the accident—performed in accordance with operational instructions. Records showed that a permanent and stable steam discharge from tank vents occurred 3 hr 40 min before the accident and finished abruptly 47 s before the explosion. During the last 47 s, no outflows from vents were visible. Also, the course of the explosion was recorded. The lifting of the roof was visible, followed by a rising fireball and flames. Evidently, the tank roof was blown off by the explosion of a flammable mixture in the tank atmosphere. The investigations after the accident showed that the vents were free and that the steam pipeline to the steam-purging inlet was open and free, too. Searching for accident causes. The fire triangle describes three requirements that have to be fulfilled for a fire/explosion of a gas mixture: an oxidant, a fuel and an ignition source.2 Accident causes combine the three requirements. Identifying the direct accident cause was not possible without identifying the specific oxidant, fuel and ignition source that were present inside the bitumen storage tank. Steam inlet pipe Vent nozzles

Bitumen circulation inlet pipe

S

FIG. 2

36

N

Bitumen storage tank; an aerial view with selected openings.

I NOVEMBER 2010 HYDROCARBON PROCESSING

Oxidant. For at least four days before the

accident, steam purging had been inactive on the tank. Its gaseous volume communicated freely with the atmosphere outside vent openings. Tank space above the liquid level undoubtedly contained mainly air at the start of steam purging 3 hr 40 min before the accident. Sweep steam purging was used to make the atmosphere of the tank inert, but it was not able to perfectly mix the whole gaseous volume when the liquid level was low.2 Steam is much lighter than air (and hydrocarbon vapors). The arrangement of the steam inlet pipe did not make the steam move into the lower parts of the tank. Only the upper part of the tank (about 20% of its total volume, according to an estimation made by plant personnel) is believed to have been filled with a steam blanket. Lower parts of the gaseous volume probably still contained mainly air during the explosion. Large volumes of bitumen in the storage tank were used from time to time as terminal volume to empty various connected pipes by air. These emptying operations brought additional oxidant into the tank. Fuel. Flammable, gaseous substances had

to be present in the atmosphere inside the tank in a concentration above the lower flammability limit for the explosion to occur. The bitumen itself releases a certain amount of light hydrocarbons, but measurements indicate that the total content above the bitumen level is one order below any conceivable lower flammable limit (LFL). The bitumen present inside the tank before the accident was of standard quality; therefore, the fuel source for the explosion had to be found elsewhere. The filling pipeline came into the tank from a manifold to which pipelines from a few other storage tanks were also connected. Analyzing operational records showed that asphalt varnish was pumped through a pipeline that was connected to the manifold, more than two days before the accident. The asphalt varnish represented a mixture of bitumen (identical with the stored one) and lacquer diluents. If a check valve in the pipeline that was used for pumping asphalt varnish had not worked properly, a certain amount of varnish would have entered the manifold. Insufficiently closing the check valve in its closing direction is a rather frequent defect that cannot be excluded. The liquid amount that would have entered the manifold in this case might have easily reached many liters. Undesir-

able liquid containing light hydrocarbons would then have been transported into the bitumen storage tank as soon as any of the pipelines connected to the manifold would have been emptied into it. Other potential sources of fuel such as catalytic cracking on steam heating pipes or steam reforming are not considered to be probable since temperatures inside the tank were not high enough. Observed explosion outcomes enabled estimating the amount of light hydrocarbons that had to be present in a flammable cloud inside the tank. The light hydrocarbons originated from lacquer diluents, the boiling interval was 135°C–220°C, LFL is 0.8% vol. and upper flammable level (UFL) is 6.5% vol. To assess the minimum amount of light hydrocarbons necessary to have lifted the tank roof off, it was necessary to start with the overpressure, which could have caused it. Overpressure at 12 kPa is enough pressure to lift a storage tank roof off.3 The question is, how large does the flammable cloud have to be if it is capable of generating 12 kPa of overpressure inside the tank? From the state equation, it follows that if the vapor space volume inside the tank is 989 m 3, then an increase in the vapor volume should be equal to 117 m3 under normal pressure. This volume increase is caused by generating hot combustion products. The number of moles inside the tank should not change during combustion. Only the temperature difference between the initial and final states could cause the volume increase. The system’s initial temperature was supposed to be equal to 478 K. The combustion products’ temperature was estimated to be 1,500 K. This temperature is in accordance that the flame temperature at the LFL for methane is 1,498 K and approximately 1,573 K for other lower paraffinic hydrocarbons.4 Comparing these final and initial states, an expansion factor equal to 3.14 was obtained. The equation for the volume of the explosive mixture capable of producing the given pressure increase is: Vexpl + 117 = Vexpl × 3.14. This results in 54.7 m3 of the explosive mixture, with a concentration equal to LFL. The light hydrocarbons may be represented by C9 fraction with a mean molecular weight of 148.4 g/mol. Using the molecular weight, the evaporated flammable vapor amount is 1.66 kg. It is certainly the lowest possible amount, not taking into account the product cooling and venting through the two vent nozzles. Higher amounts of


PLANT SAFETY AND ENVIRONMENT evaporated flammable vapors (e.g., 5 kg and 10 kg) would lead to higher values of the calculated overpressures (36.2 kPa and 72.4 kPa, respectively). Since these values are well above the 12 kPa, neither cooling nor venting through two vent nozzles would have prevented the roof from lifting off. Such an amount of flammable vapors could have easily originated in the asphalt varnish that entered the tank via the manifold. The flammable cloud could have been formed after the temperature increase during the inflow of hot bitumen between –3:40 and –0:40 hr. Hydrocarbon vapors are heavier than air so the operation of sweep steam purging would not have removed them from the tank with a low bitumen level. Fig. 3 illustrates the situation that is supposed to have been established inside the tank after adding hot bitumen. Possible ignition sources. Some igni-

tion sources may include hot work, static electricity, hot surfaces, pyrophoric iron sulfides, pressure (compression ignition), friction and mechanical sparks, sudden decom-

pression and catalysts.5 Some alternatives may be excluded immediately. There was no hot work carried out at the tank weeks before the accident. A compression or decompression sharp enough to ignite the flammable vapors is not conceivable under conditions inside the tank. No moving mechanical parts that would be able to cause friction or sparks were present inside the tank. Movement of nonconductive liquid into the tank finished at least 25 min before the accident; hence collection and discharge of static electricity are not considered to be probable to ignite the explosion. Hot surfaces, in the usual meaning of this term, were not present inside the tank. However, a layer of coke sediments were found on the south–southwestern wall, burning intensely after the explosion. Suspicion arose—the coke sediments had been smoldering even before the accident and they ignited the flammable cloud. The possible presence of catalysts (e.g., coke particles with large active surfaces) in liquid bitumen was considered, too. The presence of catalytic surfaces could

SPECIALREPORT

cause a decrease in the auto-ignition temperature of flammable vapors and lead to ignition after an induction period.2 Without a catalyst, auto-ignition of lacquer diluent vapors is not possible under 240°C. Pyrophoric iron sulfides form when iron is exposed to hydrogen sulfide (H2S), or any other compound that contains sulfur, in an oxygen-deficient atmosphere. Pyrophoric iron sulfide may form in heated bitumen storage tanks as the result of a reaction between H2S given off from the bitumen surface and iron in the form of rust on the tank roof.6 H2S was present inside the tank. Hence, the area of new bitumen circulation inlet welds seemed to fulfill all conditions for pyrophoric iron sulfide formation. Examining the facts and formulating hypotheses. Three possible

ignition sources were identified: smoldering coke, auto-ignition catalysts and pyrophoric iron sulfides. Examination of these three hypotheses with known facts is necessary, and Table 1 represents the fact/ hypothesis matrix.7

TABLE 1. Fact/hypothesis matrix. Legend: (+) compatible with hypothesis; (×) not likely

Fact or condition/hypothesis

Temperatures inside tank between 140°C and 205°C, steam blanket under roof

No anomalies (coke particles, hot spots) were found in unburned bitumen

Abrupt interruption of steam discharge 47 sec before the explosion

Tank roof fell beside tank in south– southwest direction

Light hydrocarbon vapors ignited by smoldering coke on south–southwest wall

+

+

×

×

Light hydrocarbon vapors auto-ignited after induction period

+

×

×

+

×

+

×

+

Light hydrocarbon vapors ignited by pyrophoric iron sulfides from new weld

Steam outlet

Steam inlet

Air Injection intake of water

Bitumen circulation inlet pipe

Steam blanket, 140°C

Position of new weld

Air with hydrocarbon vapors

Fresh air movement

Shrinking of steam and air

Hydrocarbon vapors raised by water evaporation Bitumen, 205°C

FIG. 3

Situation inside the tank, 30 min before the accident.

FIG. 4

Situation inside the tank during water injection.

HYDROCARBON PROCESSING NOVEMBER 2010

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PLANT SAFETY AND ENVIRONMENT

TABLE 2. Modified fact/hypothesis matrix. Legend: (+) compatible with hypothesis; (×) not likely

Fact or condition/hypothesis

Temperatures inside tank between 140°C–205°C, steam blanket under roof

No anomalies (coke particles, hot spots) were found in unburned bitumen

Abrupt interruption of steam discharge 47 s before the explosion

Roof of the tank fell beside tank in south to south-west direction

Light hydrocarbon vapors ignited by smoldering coke on south– southwest wall after the introduction of a small amount of water

+

+

+

×

Auto-ignition of light hydrocarbon vapors occurred after the introduction of a small amount of water

+

×

+

+

Light hydrocarbon vapors ignited by pyrophoric iron sulfides from new weld after introducing small amounts of water

+

+

+

+

Welding of the bitumen circulation pipe inlet— approximately 6 weeks before

Emptying of pipe line, 10 m3 of air introduced inside the tank between -0:40 and -0.25 hr

FIG. 5

Return of bitumen tank back into operation Approximately 4 weeks before

Introduction of waste into steam inlet pipe at -00:00:47 sec

Penetration of light hydrocarbons into bitumen inside tank Approximately 2 days before

Interruption of steam discharge between -00:00:47 and 00:00:00 sec

Penetration of fresh air to pyrophoric iron sulfides at new weld < 00:00:00

Start of steam purging At -03:40 hr

Spontaneous ignition of pyrophoric iron sulfides, release of sparks < 00:00:00 sec

120 t of bitumen added, level raised to 180 cm, temp. to 205°C between -3:40 and -0:40 hr

Explosion of air and hydrocarbon vapors mixture at 00:00:00 sec

Development of the accident.

Coke smoldering can develop slowly and gradually, so there would have been no reason for an abrupt interruption of steam discharge. If the explosion was ignited by smoldering coke at the south–southwestern wall, then the final position of the blown roof would have been expected on the opposite side of the tank. Catalyzed autoignition would require the presence of a catalytic surface on the bitumen. However, no corresponding anomalies were indicated in the liquid. Auto-ignition develops slowly and gradually, thus giving no explanation for the abrupt interruption of steam discharge. Pyrophoric iron sulfides spontaneously ignite after they dry out and come in contact with air, but there was no indication that the steam blanket was replaced by air at new welds. Again, the abrupt interruption of steam is not compatible with the hypothesis. Water injection. None of the consid-

ered ignition sources were able to explain satisfactorily the abrupt interruption of steam discharge 47 sec before the explosion. Evidently, the interruption did not result in any of the conceivable ignition 38

Bitumen level 65 cm, temperature 155°C— more than 1 day before

I NOVEMBER 2010 HYDROCARBON PROCESSING

processes. It resulted from an additional cause and probably contributed to the ignition process. Closing or plugging the steam pipeline or the outage of the steam supply system would have caused a slow decrease of steam discharge, not an abrupt interruption. An event must have occurred that caused an immediate pressure decrease inside the gaseous volume of the tank. Such an event could have been the steam pipeline plugging with water. Steam lines need to be equipped by steam traps. A steam trap is a device used to discharge condensate and non-condensable gases while not permitting live steam escaping. If the steam trap is not present or if it fails, then a water plug may form inside the pipeline. There are indications that a steam trap was not present in the lower part of the steam line to the steam purging inlet, so a water plug formation seemed possible. If one liter of water had penetrated into the steam purging line it would have created a water plug about 2 m long. The plug would have been transported into the end piece of the steam purging pipe. The water would have been injected into the tank’s vapor space through tiny holes in the end

piece at 4.5 bar. The injection would have abruptly cooled steam and gases inside the space, causing the gases to shrink. This results in under pressure inside the space, reversing the flow through the vent nozzles. The situation is illustrated in Fig. 4. In Fig. 4, superheated water leaves the holes in the end piece and some of it evaporates immediately (approximately 7.9%). Flash evaporation of water creates an expanding zone around the end piece inside. The temperature decreases to the boiling point of water (100°C). The volume of this zone is relatively small since one liter of water creates 0.14 m3 of flash evaporated steam. Tiny droplets of boiling water fly away from the expanding zone into much warmer steam and/or air around and below. The droplets are heated and evaporated. The tank atmosphere cools down and shrinks. Globally, 1.7 m3 of steam will emerge from 1 L of water. Simple calculations show that, in evaporation and balancing temperatures in 100 m3 of steam, the steam is cooled by 18.2°C and shrinks to 4.4 m3. Analogously, 100 m3 of air would be cooled by 24.8°C and shrunk to 6.2 m3. This process leads to a movement inside the tank atmosphere towards the expanding zone around the end piece of the steam-purging pipe. Possible evaporating droplets on the tank walls make this movement even stronger. If some droplets fell on bitumen liquid level, they would evaporate and raise hydrocarbon vapors above the liquid surface. The mixture of water aerosol and cold steam is relatively heavy and tends to sink into the air and hydrocarbon mixture. Expansion in the steam-aerosol area and turbulences caused by water evaporation make the steam-aerosol area less permeable for downward flowing gas, especially in the vicinity of the end piece.


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PLANT SAFETY AND ENVIRONMENT The estimations confirm that after water injection, the overall balance of water evaporation (with a positive influence on pressure inside the tank) and steam/air shrinkage (with a substantially higher negative influence on pressure inside the tank) will result in air intake through vent nozzles. Cold and relatively heavy fresh air tends to sink through the steam pad into the air and to dilute the air and hydrocarbon mixture. However, there was a zone consisting of steam and aerosol that was denser than steam and obstructive due to turbulences. This, combined with the sinking and shrinking movements inside the zone, would have caused part of the inflowing air to be sucked in a perpendicular direction between the zone and the tank roof and to move along the roof into the peripheral parts of the tank (Fig. 4). Sinking fresh air and the steam pad at the southern wall may have also caused air movement at the opposite side upward against the wall.

into the bitumen inside the tank; the creation of pyrophoric iron sulfides at the new weld; and the introduction of water into the steam inlet pipe. The accident would never have occurred if a small amount of water had not been introduced into the steam purging inlet pipe. It is highly probable that similar situations had occurred in the tank’s 40-yr history, but the necessary causes never coincided. Fig. 5 illustrates a probable multiple-root cause analysis, which may lead to some lessons learned from the accident while recommending proper corrective measures. HP 1

2

3

4

Possible ignition sources after water injection. Three new hypotheses

on possible ignition sources represent the interactions of sources that were selected with water injection. The results of the examination are summarized in Table 2. If the source were smoldering coke on the south–southwest wall after the water injection, then the abrupt interruption of the steam discharge is understandable. However, there is no new explanation of why the tank roof fell in the opposite direction than would be expected. Similarly, for catalytic auto-ignition the presence of a catalyst remains unexplained. The situation has changed only for pyrophoric iron sulfides. The preceding section explains that fresh air could penetrate to a new weld after water injection. Hence, conditions for spontaneous ignition of pyrophoric iron sulfides would have been fulfilled. There is evidence that such an ignition may produce sparks.6 The immediate contact between the sulfide and the flammable mixture is not necessary. Consequently, the pyrophoric iron sulfides at the new weld might have acted as an ignition source after introducing small amounts of water into the steam inlet pipe. Conclusions. The analysis helped com-

plete the time line of events leading to the accident (Fig. 5). Dashed borderlines denote the events that cannot be proved. Probably, the combination of four direct causes led to the accident: the inefficient sweep steam purging of the bitumen storage tank; the penetration of light hydrocarbons

5

6

7

LITERATURE CITED Kletz, T. A., New Fire-fighting Agent Meets Opposition, By accident ... a life preventing them in industry, PFV Publications, London, 2000 Crowl, D. A., “Understanding Explosions,” American Institute of Chemical Engineers, New York, 2003. Kletz, T. A., “Myths of the Chemical Industry,” The Institution of Chemical Engineers, Rugby, 1985. Lees, F. P., Loss Prevention in the Process Industries, Second Edition, ButterworthHeinemann, Oxford, 1996. IRP18 Committee, University of Calgary, Canada, http://www.firesandexplosions.ca/hazards/ ignition sources.php, September 24, 2007. Davie, F. M., T. W. Nolan and S. Hoban, “Study of Iron Sulfide as a Possible Ignition Source in the Storage of Heated Bitumen,” Journal of Loss Prevention in the Process Industry, Vol. 6, Issue 3, pp. 139–143, September 1993. Center for Chemical Process Safety, “Guidelines for Investigating Chemical Process Incidents,” American Institute of Chemical Engineers, New York, 1992.

Milos Ferjencik studied nuclear engineering at Prague Technical University and graduated in 1981. He worked at Nuclear Research Institute, and in various technical and research positions. In 1991, Mr. Ferjencik concentrated on chemical risk analysis. In 1995 he started his own consultancy profession. Mr. Ferjencik was an independent risk and reliability consultant and an external teacher at University of Pardubice. Most recently, Mr. Ferjencik has worked as a full-time assistant professor of safety engineering at the University of Pardubice.

Bretislav Janovsky received BS and PhD degrees from the University of Pardubice focusing on theory and technology of explosives. He worked in various technical and research positions at the University of Pardubice and a privately held company, TLP Prague. Dr. Janovsky concentrated on consequence analysis under risk analysis in the process industries. He was a full-time assistant and associated professor of safety engineering at the University of Pardubice and started his own professional consultant business in 1995. Today, Dr. Janovsky is the research and development director at OZM Research Bliznovice and teaches part-time at the University of Pardubice.


PLANT SAFETY AND ENVIRONMENT

SPECIALREPORT

Consider new analysis for flares Applying dynamic models in designing safety systems can reduce capital costs Z. URBAN, M. MATZOPOULOS and J. MARRIOTT, Process Systems Enterprise Ltd., London,UK, and B. MARSHALL, Softbits Consultants Ltd., Medstead, UK

A

pplication of dynamic modeling for relief system design can substantially lower capital expenditure (CAPEX) while simultaneously improving plant safety. This article considers using dynamic analysis to two areas: vessel depressurization (or “blowdown”) and flare network design. New modeling methods can accurately quantify relief loads and metal temperatures to enable informed safety and CAPEX decision support. CASE HISTORY 1: VESSEL DEPRESSURIZATION

Detailed dynamic analysis of the rapid depressurization (blowdown) of high-pressure vessels is a key element of the safety analysis of oil and gas facilities and other high-pressure installations. Event description. Depressurization of a vessel usually results in cold gas venting into the flare system. The cold gas can significantly lower the temperatures within the process equipment metal walls and pipework, as well as the relief system pipework immediately downstream of the blowdown valves (BDVs). Low temperatures can lead to embrittlement of the equipment and pipework metal walls, and the difference in temperature between adjacent metal sections can result in high thermal stresses. This condition has implications for the integrity of process vessels, pipework and sections of the relief system, as well as for CAPEX. Accurate analysis of likely relief scenarios is essential to determine: • Relief loads entering the flare network. For new designs, accurate information is needed to achieve an optimal design that minimizes the piping diameters required to meet Mach number and back-pressure constraints. Minimizing the piping sizes also provides benefits in terms of reduced support infrastructure, which

can be particularly important in the case of offshore platforms where additional weight is heavily penalized. For revamps or expansions to an existing process plant, accurate data can help determine whether the current flare system can handle the new loads acceptably. In either design scenario, CAPEX savings can be considerable. • Temperature throughout the process and pipework metal walls to identify areas of potential embrittlement, and where (and when) unacceptable thermal stresses are likely to arise. Such information can be used to mitigate potential problems either by controlling the relief rates or by rerouting the relief flows. • Temperature of the relieving “gas” streams (which may actually contain evaporating entrained liquids). This provides essential information for choosing the appropriate material of construction for the critical sections of pipework immediately downstream of the BDV. The effects of low temperature can usually be addressed by using suitable materials of construction. Unfortunately, in some cases, such materials can be expensive, and it is highly desirable to minimize their use without compromising safety considerations. This requires accurate quantification of flowrates and temperatures of the relieving stream, as well as the minimum temperatures reached in the metal walls. Complex phenomena. Depressurization of a vessel involves a complex set of coupled physical phenomena that must be characterized accurately to understand behavior and provide suitable design values. Current depressurization modeling is often performed with off-the-shelf process flowsheeting simulators that use an equilibrium thermodynamic approach. The latter provides some indication of the flow and temperature, but by no means adequately

describe the complex thermodynamic and kinetic phenomena occurring as a result of rapid decreases in pressure. The fact that multiple phases can form within the vessel, and that these may not be in equilibrium with each other or with the vessel wall, can have a significant effect on both the relief flows and metal temperatures of the vessel and relief system pipework. The sudden decrease in pressure in a gas-filled vessel results in a rapid change in the thermodynamic state of the gas within the vessel. This can result in nucleation of liquids within the gas bulk to form a “droplet phase” as shown in Fig. 1. Some of the nucleated liquid leaves as entrained droplets in the high-velocity gas exit stream (Fig. 2). Downstream of the vessel, and as the pressure further reduces, this exiting entrained liquid evaporates into the bulk gas stream, lowering the temperature of the cold exiting stream even further. This creates a risk of brittle fracture of the flare system pipework. Some proportion of the liquid remaining in the vessel drops to the vessel floor. Gas to relief system

Nucleating liquid droplets

FIG. 1

Formation of droplet phase in blowdown event.

HYDROCARBON PROCESSING NOVEMBER 2010

I 41


SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

Initially, this evaporates instantly due to the warm temperature of the metal it encounters. The effect is similar to a drop of water falling on a hotplate. However, once the metal has cooled sufficiently (typically after a few tens of seconds), liquid begins pooling (Fig. 2) and forming a continuous liquid phase.

Entrained liquid droplets leaving with gas

Condensed liquid droplets pooling and evaporating FIG. 2

Droplets exiting in the gas stream and forming a continuous liquid phase on the vessel bottom.

Singlephase (gas)

The pool boils vigorously, cooling and reducing in size, and, in turn, reducing the temperature of the metal beneath it. This event can lead to significant temperature differences between the metal immediately below the pool and its surroundings—presenting a very real threat of brittle fracture and rupture of the vessel base. The effect of the phenomena can be seen graphically in Fig. 3, which shows the results of depressuring a vessel filled with light-hydrocarbon gas at 120 bar. In the initial phase of depressurization, the gas temperature (black line) drops rapidly. The temperature of the metal wall in contact with the gas (green line) begins to drop, but much more slowly because of the resistance to heat transfer between between wall and gas and heat conduction within the wall. After about 80 seconds, a droplet phase begins to form throughout the gas. Initially, droplets in contact with the metal heat up rapidly and vaporize (red spike). When cool liquid droplets (at a temperature close to that of the bulk gas) begin to pool on the vessel floor, the liquid temperature increases further above the bulk gas temperature as the liquid is heated by the metal wall and changes in composition. After a while, the gas bulk temperature begins to rise because of heat influx from the metal wall.

Two-phase (gas, droplet) Three-phase (gas, droplet, liquid)

CASE HISTORY 2: FLARE NETWORK

Onset of nucleation

Formation of liquid phase and pooling

300 Metal wall in contact with gas

290 Temperature, K

280 Equilibrium

Metal wall in contact with liquid

270 260 250

Liquid Bulk gas

240 230 0

FIG. 3

42

100

200

300

400

500

600 700 Time, sec

800

900

Temperature profiles over the duration of the blowdown event.

I NOVEMBER 2010 HYDROCARBON PROCESSING

The items of most concern are the rapidly decreasing temperature of the metal in contact with the liquid pool (blue line) and the difference between the temperature of this metal and the adjacent metal contacting with the gas (green line). The metal temperature can be seen to drop to nearly –30°C, approaching the brittle fracture temperature for carbon steel. The temperature difference between vessel floor and sides rapidly increases to over 20° and is nearly 40° by the end of the blowdown, which may give rise to unacceptable stress. Because of the rapid change of conditions, the three phases coexisting in the vessel (gas, droplet and a pool of liquid,) and the vessel walls are not in equilibrium with each other throughout most of the blowdown event. In comparison, the dotted line shows the equivalent temperature curve obtained using an equilibrium model for the same blowdown, which predicts a much less severe drop in temperature. This model fails to identify the most significant safety-related aspect—the cooling effect of the liquid on the vessel bottom. Fig. 4 shows the resulting vessel wall temperatures and associated thermal stresses for the vessel vertical walls as color temperature plots. This information would not be available without rigorous modeling of the nonequilibrium mass and energy transfer between phases. This example describes just one scenario. Other scenarios may develop depending on the initial inventory and state of the material in the vessel. For example, there may be “bubblet” nucleation in super-critical fluid, rather than the droplet nucleation described here.

1,000 1,100 1,200

Conventional flare header design techniques use peak relief flows in steady-state simulation to assess system capacities and determine back-pressures downstream of blowdown valves (BDVs) and pressure safety valves (PSVs), Mach number in the headers, and radiation at the flare tip. This steady-state assumption is highly conservative. While conservative approaches may be desirable in safety system design, they can nevertheless lead to gross overdesign throughout the system. Key areas of over design include: Oversized flare header. Sizing the header for the sum of the maximum flows takes no account of effects such as: • System packing, where the gas pressurizes the available volume in the flare network


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PLANT SAFETY AND ENVIRONMENT

• Potential for sequencing of flare events. For example, depressurization initiated deliberately by an operator may be complete well before a fire causes PSVs to lift. Steady-state peak flow analysis, on the other hand, assumes that all events occur simultaneously. Reducing the peak flows used as the design basis by judicious analysis can significantly reduce pipe sizes and materials and fabrication costs, which can be substantial for large-diameter headers. Reducing the size also creates knock-on savings related to the support structure and flare stack size. Oversized flare stack. The flare stack

sizing depends on radiation emitted by the flame, which is a function of the volumetric gas flowrate through the flare tip. Using unrealistically high flowrates determined from peak flows results in an over-long stack, creating weight problems in offshore facilities or adding stack support costs (or unnecessary additional header length) in onshore facilities. Similarly, a lack of accurate temperature information leads to a wide span between the minimum and maximum design temperatures used for gas arriving at the stack, resulting in unrealistic allowances for thermal expansion and contraction. Over-use of expensive alloys.

Although flare system pipework may be in contact with gas at extremely low temperatures, this typically occurs for a relatively short duration. The use of steady-state flows does not consider the duration of such exposures to low temperature, which may result in very conservative and expensive application of alloys. It can be argued that a good flare net-

FIG. 4

44

work design is one that minimizes capital expenditure while meeting all safety constraints. Overdesign should be avoided wherever possible. By making simple dynamic analyses using data that is mostly already available in some form, it is often possible to refine network designs to arrive at systems with a significantly lower capital cost while demonstrably meeting safety requirements. Similarly, it is often possible to find additional capacity during retrofits, thus removing the need for additional capital expenditures. Typical examples of where dynamic analysis can bring significant new information that has an impact on capital cost are: Peak flowrates. The actual relief flow through any PSV is at the maximum only for a short period. Using steady-state methods based on peak flows is equivalent to making the assumption that all relief flows start at the same time and go on forever. In reality, it is often possible to take credit for staged or staggered relief. Shifting depressurization of certain units by a few tens of seconds can make a significant difference to the peak flows through the system—an effect that cannot be represented at all by steady-state simulation. Packing. Steady-state approaches make the implicit assumption that the flare system has no volume—what goes in comes out, instantly. For larger systems, the impact of relief flows is partially “absorbed” by pressurization of the flowing lines and the dead volumes in non-flowing parts of the system. This “packing” effect can reduce both the calculated peak back-pressures or Mach numbers and the peak flows seen at the flare tip, allowing reduction in header

4a. Wall temperatures at the end of the blowdown; 4b. Wall thermal stresses from the effects of pooling liquid.

I NOVEMBER 2010 HYDROCARBON PROCESSING

FIG. 5

and tailpipe diameters and flare stack lengths, respectively. Dynamic simulation allows this important buffering effect to be taken into account in the design. Duration. Equally important, dynamic simulation can be used to determine the duration of peak flare loads. Engineering judgment can then be used to assess the risks of any infringements. For example, a 5-second violation of back-pressure or radiation constraints may well be acceptable, especially given the capital costs of oversizing the flare system to avoid such a contingency. Temperature. Relief system pipework that is likely to encounter low temperatures needs to be constructed from expensive alloys such as Inconel to avoid the possibility of embrittlement and consequent fracture. The true extent of pipework that truly needs to be constructed of such materials is impossible to gauge with steady-state simulation, as low-temperature flows are considered to continue forever, ensuring that calculated metal temperatures reach their minimum. In reality, such flows may only last for a few minutes; the thermal inertia of the pipework metal and heat gain from the environment prevent the pipework from reaching the gas temperatures during this time (a similar effect can be seen in the bulk gas and metal temperature plots in Fig. 3). It is frequently possible to reduce the usage of alloy significantly based on the more accurate information from the dynamic analysis. One oil company reported saving $1.5 million on a single vessel this way. Flare-stack temperatures. Dynamics can also help provide a true picture of the temperature of gas arriving at the flare tip. Proper calculation of the effect of low-tem-

Example of a flare network showing active sources.


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SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

0.30

0.5

Mach number

Mach number

0.4 0.20

0.10

0.3 0.2 0.1 0.0

0.00 0

FIG. 6

100

200

300

400 500 Time, sec

600

800

Mach numbers of 24-in. header in a flare system.

TABLE 1. Back-pressures for 24-in. header case Inlet source

Back-pressure, bar

Train_1_DP

1.63

Train_2_DP

1.75

Train_3_DP

1.63

Train_4_DP

1.79

Train_5_DP

1.68

TABLE 2. Back-pressures for 20-in. header case Inlet source

Back-pressure, bar

Train_1_DP

2.23

Train_2_DP

2.32

Train_3_DP

2.23

Train_4_DP

2.34

Train_5_DP

2.27

perature gas over a short time taking into account metal-thermal inertia and ambient heating provides much more accurate minimum and maximum design temperatures, allowing the designer to make sensible decisions on stack length and support mechanisms. Example. Consider the flare system

shown in Fig. 5, where the header sizes are set primarily by a depressuring scenario from five units simultaneously, as highlighted in the figure. One of the key questions is the size of the long main-header pipe leading to the flare stack. Typically, such a system is designed by working back from the flare tip, sizing all the lines based on velocity constraints until reaching the relief valves, and then confirming that other constraints, such as back-pressure constraints at the relief valves and limits on noise, are not violated. In this study, a system has been designed using steady-state techniques using a veloc46

700

I NOVEMBER 2010 HYDROCARBON PROCESSING

900

0

FIG. 7

100

200

300

400 500 Time, sec

600

700

800

900

Mach numbers of 20-in. header in a flare system using a staggered blowdown.

ity heuristic requiring a Mach number between 0.25 and 0.35 in the main lines. A conventional steady-state approach calculates the Mach number using the sum of the peak flowrates. The maximum Mach number at the pipe outlet is represented by the dotted line in Fig. 6. At 0.29, this is well within the 0.25–0.35 range. The back-pressures at the five blowdown valve sources are listed in Table 1; these are well below the limit. For illustration, it is useful to do a “pseudo-dynamic” run, using relief flow curves but taking no account of the flare system volume. This shows a Mach number profile over time that has the characteristic sharp-peaked shape of relief flow curves (red line in Fig. 6). As expected, the peak Mach number from this run is the same as for the steady-state case, at the sum of the individual peak flows. Although this case adds no new information to the design, it does provide some indication of the length of the blowdown event, allowing judgment to be applied in the case of constraint violation. If a full dynamic simulation is done, taking the volume of the system (both for active and inactive branches) into account, it can be seen that the effect of flare system packing significantly reduces the peak Mach number observed, to about 0.25. It is evident from these results that there may be potential to reduce the diameter of the 24-in. header, as the Mach number is nowhere near its limit. A new series of calculations is performed with a 20-in. header diameter to see the effect of reducing the flare system line sizes. As expected, the Mach number obtained using steady-state peak flows (0.4) violates the system design constraints, indicating that the design is not viable (Fig. 7, dotted line). The corresponding pseudo-dynamic case shows (Fig. 7, red line) that the value

is out of range for about three minutes, which is also unacceptable. However, a full dynamic simulation taking into account line packing shows that the peak Mach number is within the 0.35 limit (Fig. 7, black line); the back-pressures (Table 2) remain well within the limits. The added information provided by the dynamic simulation thus indicates that the design is indeed viable. If further mitigation is required, it is possible to investigate the dynamic effects of staggering the depressurization, so that units depressure in sequence. The green line in Fig. 7 shows the effect on Mach number of delaying the blowdown of Unit 2 by several minutes. Similar approaches can be applied to retrofit cases, often demonstrating that it is possible to accommodate additional sources in an existing flare system that is ostensibly operating close to its limits. Conclusion. The dynamic simulation capabilities of modern software tools provide a number of options for analyzing both the depressurization event—to determine accurate relief flows and fluid thermodynamic conditions—and the flare header design itself. This enables engineers to design systems that comply with safety guidelines based on a much more realistic representation of behavior than traditional methods allow, and, at the same time, to identify opportunities for significant capital savings. In the case of depressurization, rigorous dynamic simulation identifies potentially dangerous situations. For new flare system designs, it can lead—among other benefits— to a reduction in header size, resulting in significant capital savings. For existing headers, it provides a means to establish whether there is sufficient capacity to accommodate new sources, thereby avoiding the need for a new header and flare. HP


PLANT SAFETY BIBLIOGRAPHY Haque, M. A., S. M. Richardson, and G. Saville, “Blowdown of Pressure Vessels. I—Computer Model,” Transactions of the Institute of Chemical Engineers Part B: Process Safety Environmental Protection, 70(BI), 1, 1992. Mafgerefteh, H. and S. M. A. Wong, “A numerical blowdown simulation incorporating cubic equations of state,” Computer Chemical Engineering, Vol. 23, p. 1309, 1999. Gruber, D., D.-U. Leipnitz, P. Sethuraman, M. A. Alos, J. M. Nougues and M. Brodkorb, “Are there alternatives to an expensive overhaul of a bottlenecked flare system?” Petroleum Technology Quarterly Q1, 2010. Chen, F. F. K., R. A. Jentz and D. G. Williams, “Flare System Design: A Case for Dynamic Simulation,” Offshore Technology Conference, May 4–7, 1992, Houston. Goyal, R. K. and E. G. Al-Ansari, “Emergency Shutdown devices and relief system sizing and design in oil refineries,” Hydrocarbon World, Vol. 4, No. 1, Touch Briefings, 2009. Speranza, A. and A. Terenzi, Blowdown of Hydrocarbons pressure vessel with partial phase separation, Series of Advances in Mathematics, available from http://www.i2t3.unifi.it/upload/file/Articoli/ animp_2004.pdf, 2005. Szczepanski, R., “Simulation programs for blowdown of pressure vessels,” IChemE SONG Meeting, 1994.

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Designing the correct pressure-relieving system Use these relief rate calculations for gas thermal expansion as a cause for overpressure S. RAHIMI MOFRAD, Petrofac Engineering and Construction, Sharjah, UAE

T

he first step in designing a pressure-relieving system is to specify all possible causes of overpressure that is applicable to different equipment. The American Petroleum Institute (API) standard 5211 introduces the various emergency cases and general guidelines for calculating relief rates. While liquid thermal expansion is widely recognized as causing overpressure, gas thermal expansion is rarely taken as a credible cause of overpressure. Standard guidelines may be absent or it may be believed that gas thermal expansion has negligible contribution to overpressure. As a general practice for liquid thermal expansion, thermal relief valves (TRVs) are normally provided for equipment operating full of liquid. These can be blocked-in between inlet and outlet valves, where sufficient heat may be applied to fluid, increasing the pressure above the equipment design pressure.2 The relief requirement of a vessel under fire containing only gas (unwetted vessel) was developed by API 521 and other reference materials.3 This article will present cases where gas thermal expansion is caused by heat sources other than external fire—including process hot stream, solar radiation and ambient temperature variation. Gas expansion fundamentals. Gas thermal expansion occurs in all equipment and piping with the following conditions:

• Line or equipment is isolated for operational or emergency purposes • A heat source with temperature higher than gas temperature exists. For the described system, a relief valve is needed if: • Pressure increased since the temperature increased, exceeding the system design pressure • No other overpressure protection device on the system, no provision to prevent gas tight blockage such as a locked open (LO) valve, leaking check valve or venting procedure before closing isolation valves.2 Figs. 1A and 1B show the pressure rise of blocked air, based on the results of the Soave–Redlich–Kwong (SRK)4 equation of state (EOS) for different pressures and temperatures. As shown in Fig. 1A, the blocked-in gas pressure rise due to temperature changes (ΔP/ΔT), is not significant in low to medium pressure applications. However, for high pressure gases it can reach 2 bar/°C (Fig. 1B). The values given in Figs. 1A and 1B may be used to estimate the potential pressure rise for other gases, but for gas mixtures with a wide composition range, use the suitable EOS. Figs. 1A and 1B illustrate that gas thermal expansion should be considered as causing overpressure in the following systems: 2.5

0.10 1 bara 5 bara 10 bara 20 bara 30 bara

0.09 0.08

2.0

ΔP/ΔT, bar/1°C

ΔP/ΔT, bar/1°C

0.07

100 bara 200 bara 300 bara 500 bara

0.06 0.05 0.04

1.5

1.0

0.03 0.5

0.02 0.01 0.00 -50

0

50

150 100 Temperature, °C

200

FIG. 1A Gas expansion for low to medium pressure.

250

300

0.00 -50

FIG. 1B

0

50

150 100 Temperature, °C

200

250

300

Gas expansion for medium to high pressure.

HYDROCARBON PROCESSING NOVEMBER 2010

I 49


SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

• Systems with high operating pressure where ΔP/ΔT is significant. This means that even small temperature changes may cause catastrophic piping or equipment failure containing dense gas. Example, a 300-bar air cylinder stored at 23°C, the pressure will increase 1.3 bar per each degree centigrade, as shown in Fig. 1B. Therefore, if the air condition fails and room temperature increases to 40°C, cylinder pressure can reach 322 bar. Another example is a high-pressure pipeline with a possible blockage, along with exposure to sun radiation. • System with high differential temperature between heat source and blocked gas. A 20-bar cold fuel gas is superheated in a gas–gas heat exchanger using hot flue gas at 400°C. If the fuel gas is blocked, it can reach the hot gas operating temperature. Although the gas expansion is small at 20 bar, the high differential temperature can easily increase the blocked gas to pressure beyond the design pressure. Relief rate calculation. There is not a common approach

for gas thermal expansion relief studies that is applicable to all systems. These systems should be reviewed case by case. What is common among different systems is that pressure can be maintained at a safe level if the excess mass is released. Writing a mass balance equation for blocked gas gives the required relief rate:

dm(t ) (1) dt Assuming: • Since the gas composition does not change during relief, gas molecular weight is constant. Hence, m(t) = MW n(t) • Enough heat is supplied to keep the blocked gas at relieving pressure during relief (independent of time (P = PR )) • n = PV/RT predicts gas thermodynamic behavior. Eq. 1 could be rewritten as: W (t )=

W (t ) =

PR V MW d (1 /T ) dt R

(2)

Rearranging Eq. 2 results in Eq. 3, and can be solved if the temperature variation with time is known. P V MW dT W (t ) = R (3) dt RT 2 Initial checking. An appropriate safeguard for gas thermal expansion is equipment design pressure. Before starting the relief rate calculation, it should be checked at which temperature the design pressure will be reached using Eq. 4.

0.84

Ethane (mole %)

0.08

Propane (mole %)

0.04

i-butane (mole %)

0.02

MW

50

Value

Methane (mole %)

n-butane (mole %)

0.02 19.97

PN (barg)

10

PD (barg)

12

I NOVEMBER 2010 HYDROCARBON PROCESSING

(4)

A simplified form of the ideal gas law, T = TN,ave PD /PN, considering its limitations and applications can also be used. If the calculated temperature, T, is higher than the heat source temperature, no pressure safety valve is needed. In other words, the pressure of blocked gas does not reach relieving pressure even if the temperature increases from the initial temperature to the heat source temperature. If not, there is a potential for the blocked gas pressure to exceed the design pressure. The calculated temperature is the temperature where a relief valve opens for the first time (the initial relieving temperature). Relief rate calculation procedure. The following section

introduces a very simple procedure to define the blocked-gas temperature as a function of time and solves Eq. 3 for a gas-gas heat exchanger using a numerical solution. This method may result in 10%–20% overdesign on flow (required area). It’s also acceptable compared to the overdesign associated with the relief valve selection procedure—selecting a pressure safety valve (PSV) among standard orifices usually much larger than the required area. Step 1. Assume constant temperature intervals, ΔT, between the initial relieving temperature and the heat source temperature and calculate T(t+Δt):

T (t + t ) =T (t ) + T

(5)

Step 2. Calculate the average heat transfer rate between heat source and trapped gas by using Eq. 6: Q(t ) =U (t ) A (TS T (t ))

(6)

Step 3. Calculate the heat required to increase the temperature of trapped gas from T(t) to T(t+Δt) by using Eq. 7, which ignores part of the heat that is consumed by the gas container (heat-exchanger metal). q(t ) = n (t ) MWC (t ) T (t + t ) T (t ) P

(7)

Step 4. Calculate the time required to increase the temperature of trapped gas from T(t) to T(t+Δt) using Eq. 8. t =

q (t ) Q (t )

(8)

Step 5. Calculate the relief rate using Eq. 9.

W (t ) =

PR V MW (T (t + t ) T (t )) t RT (t )2

(9)

Step 6. Calculate the new number of trapped gas moles using Eq. 10 for the next calculation stage. Go back to Step 1.

TABLE 1. Gas condition and composition Parameter

PD  PN = Fig. 1value at PN andTN ,ave T TN ,ave

n (t + t ) = n (t ) W (t ) t

(10)

Case study. Table 1 shows natural gas with conditions and compositions that are superheated from 100°C–150°C in a gasgas heat exchanger using 400°C high-pressure steam. The heatexchanger tube inside diameter, total heat transfer area and total tube volume are 1 in., 100 m2 and 1.5 m3 respectively. According to Fig. 1A, the pressure rise of trapped gas at an average temperature of 125°C and 10 barg is 0.0289 bar/°C. Therefore, the minimum heat source temperature to increase the blocked-gas pressure from operating to design pressure is:


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TABLE 2. Relief rate calculation T (°C) Cp* (kJ/kgK)

␮* (cP)

␳* (kg/m3) k* (J/s m K)

Gr (—)

Pr (—)

Nu (—) U (J/sm2K)

Q (J/s)

m (kg)

q (J)

⌬t (s)

W (kg/hr)

235.7

2.84

1.75E-02

6.74

6.42E-02

7.32E+06

0.775

11.7

30.0

493205

10.10

470582

0.95

1241.1

252.1

2.90

1.79E-02

6.53

6.70E-02

5.71E+06

0.776

10.8

29.1

429724

9.78

464857

1.08

1027.3

268.5

2.96

1.83E-02

6.32

6.98E-02

4.42E+06

0.777

10.0

28.0

368505

9.47

462289

1.25

838.1

285.0

3.02

1.87E-02

6.13

7.27E-02

3.38E+06

0.778

9.2

26.9

309314

9.17

454386

1.47

669.9

301.4

3.08

1.91E-02

5.95

7.56E-02

2.54E+06

0.779

8.5

25.7

253048

8.90

449576

1.78

522.8

317.8

3.14

1.95E-02

5.78

7.86E-02

1.87E+06

0.780

7.7

24.3

199593

8.64

447697

2.24

393.8

334.3

3.20

1.99E-02

5.62

8.15E-02

1.32E+06

0.781

7.0

22.7

149015

8.40

440545

2.96

281.0

350.7

3.26

2.03E-02

5.47

8.45E-02

8.82E+05

0.782

6.1

20.8

102460

8.17

436291

4.26

185.0

367.1

3.32

2.06E-02

5.33

8.75E-02

5.25E+05

0.782

5.2

18.4

60452

7.95

434803

7.19

104.6

383.6

3.37

2.10E-02

5.19

9.06E-02

2.34E+05

0.783

4.1

14.9

24373

7.74

428127

17.6

40.5

400.0

3.43

2.14E-02

5.07

9.36E-02

0.00E+00

0.783

0.0

0.0

0

7.54

0.0

*Input from simulation software

PD PN = 0.0289 bar / °C T 125°C

Heat transfer background

Ignoring the effect of heat radiation, overall heat transfer coefficient for the tube side of the heat exchanger is calculated from the following relation:

1 1 1 1 = + + U hi kw  Aw   Ao  ho   x  Ai   Ai 

kw h x o

(12)

Substituting Eq. 12 in Eq. 11 gives U≈hi. There are many correlations for calculating the free convection heat transfer coefficient, hi, inside enclosed space. The following relation can be used for estimating this parameter when 6 x 106 < Gr Pr < 108.5 Nu = 0.104 Gr0.305 Pr0.389 where:

Gr =

Nu =

52

g (TS T ) D 3 2

T =

12 10 +125 =194.2°C 0.0289

(11)

From the heat transfer point of view, the main difference between the operating and blocked-in conditions is that the heat transfer mechanism inside the tube changes from force convection to free convection. The following assumption is applicable to blocked-in conditions:

hi 

Then:

Pr =

hi D k

I NOVEMBER 2010 HYDROCARBON PROCESSING

=

μC p k μ

Since the temperature of HP steam is higher than 194.2°C, it will over-pressurize the cold side and a relief valve is required for gas thermal expansion if the system is blocked-in. Simulation software was used to obtain the physical properties of the gas at each temperature interval. Considering relieving pressure of 13.2 barg (corresponding to a relieving temperature of 235.7°C), the relief rate was calculated (Table 2). In this example, the following points were observed: • As the blocked-gas temperature increases, the differential temperature decreases. This results in reducing the overall heat transfer coefficient and, subsequently, the relief rate. • The maximum relief load takes place at the first interval when the heat transfer rate is high. • The time between two subsequent PSV openings (Δt) is initially so short that it can be assumed as continuous relieving. The PSV will chatter as trapped gas temperature approaches heat source temperature. • Unlike liquid thermal expansion where a ¾ in. x 1 in. thermal relief valve will normally be sufficient, preliminary PSV sizing showed that a 1½ in. x 2 in. PSV with an F designation is needed for this case. • When the gas temperature reaches 194.2°C, the relief valve opens for the first time. If the heat transfer rate is high, the relief valve will remain open, otherwise, it will close until the pressure buildup is sufficient to reopen the relief valve again. The typical trends of blocked-gas parameters for this case are illustrated in Fig. 2 (Case 1). In case of low heat flux, the relief valve opens and closes repeatedly to release the excess pressure, as shown in Fig. 2 (Case 2). Gas thermal expansion can cause system over-pressurization in particular conditions that were discussed. Ignoring this case may result in a system mechanical failure. A ¾ in. x 1 in thermal relief valve recommended by API-521 may not be suffi-


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SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

cient for this case. Therefore, the relief rate has to be calculated according to the system dimension and heat transfer rate from the heat source to the blocked-in gas while an adequate PSV size needs to be utilized. In absence of any standard addressing this case, different systems should be reviewed on a case-bycase basis. HP 1 2 3 4 5

LITERATURE CITED “Pressure-relieving and depressuring systems,” API RP 521, 5th edition, January 2007. Norouzi, S. and S. Rahimi Mofrad, “What you should know about liquid thermal expansion,” Hydrocarbon Processing, November 2008. Rahimi Mofrad, S. and S. Norouzi, “Designing for pressure release due to a fire—Part 1,” Hydrocarbon Processing, November 2007. Smith, J. M. and H. C. Van Ness, “Introduction to Chemical Engineering Thermodynamics,” Fourth Edition, McGraw Hill. Cheremisinoff, N. P., Heat Transfer Handbook, Gulf Publishing Company, Houston, Texas, 2003.

Case 1: High heat flux Case 2: Low heat flux Blocked gas parameter

T=Ts Relief rate, Case 1 Relief rate, Case 2 Temperature Pressure, Case 1 Pressure, Case 2 PR

TN

NOMENCLATURE A Heat exchanger surface area, m2 Cp Gas heat capacity, kJ/kg K D Tube diameter, m g Accleration of gravity, 9.81 m/s2 Gr Grashof number, dimensionless h Heat transfer coefficient, J/s m2 K k Thermal conductivity, J/s m K m Mass of trapped gas, kg MW Gas molecular weight, kg/kgmol n Mole of trapped gas, kgmol Nu Nusselt number, dimensionless P Pressure, bara Pr Prandtl number, dimensionless Q Total heat transfer rate, J/sec q Heat content of trapped gas, J R Gas constant, 8,314 bara m3/kgmol K Ra Rayleigh number, dimensionless TS Heat source temperature, K T Blocked-gas temperature, K t Time, sec U Overall heat transfer coefficient, J/s m2 K V Trapped gas volume, m3 W Relief rate, kg/hr Z Gas compressibility factor ␤ Cubical expansion coefficient, 1/K ␮ Dynamic viscosity, Centipoises (Cp) ␷ Kinematic viscosity, m/s2 ⌬x Wall thickness, m ␳ Density, kg/m3 Subscripts i Inside o Outside ave Average w Wall N Normal operating condition D Design condition R Relieving condition

PN Time, t Gas is blocked in FIG. 2

Saeid Rahimi Mofrad is a process engineer with Petrofac

PSV opens

Typical trend of blocked-gas parameters.

Engineering and Construction at Al Soor, Sharjah, UAE. He is interested in relief load calculation, overpressure protection systems design and flare network sizing. Mr. Rahimi Mofrad has an MS degree in chemical engineering from the Sharif University of Technology and a BS degree from Shiraz University, Iran.

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SPECIALREPORT

Customize operator training for your thermal oxidizers This case history shows the benefits of site-specific programs in new equipment installations T. GILDER, Shintech Louisiana, LLC, Plaquemine, Louisiana; and D. CAMPBELL, T. ROBERTSON and C. BAUKAL, John Zink Co. LLC, Tulsa, Oklahoma

S

hintech began construction of its new Plaquemine, Louisiana, manufacturing facility (see Fig. 1) in October 2005. The facility is located on a 1,725-ac site, and it manufactures chlorine, caustic soda and vinyl chloride monomer (VCM). Historically, Shintech has manufactured only polyvinyl chloride (PVC). This new plant is Shintech’s first integrated complex. This new manufacturing facility uses state-of-the-art environmental technologies and is subject to the most stringent environmental controls in the country. Thermal oxidizers (TOs) are commonly used to treat volatile organic compounds (VOCs) and carbon monoxide (CO) emissions because TOs have very high destruction and removal efficiencies.1,2 Thermal oxidation can be defined as “the process of oxidizing combustible materials by raising their temperature above the auto-ignition point in the presence of oxygen and maintaining it at high temperature for sufficient time to complete combustion to carbon dioxide and water.”3 Shintech partnered with the manufacturer of the TOs to offer customized training to 37 of its plant operators for the start-up of its new facility.a Fig. 2 shows part of the thermal oxidization system installed at the Shintech facility. The TO training course was offered with optional Continuing Education Unit (CEUs) credits that were available to any students meeting the following criteria: take (not pass) a pre-test, attend at least 80% of the course contact time, pass (at least 80%) a post-test and complete an anonymous course evaluation. The manufacturer also operates a training organization that is accredited; the JZI b is authorized to offer CEUs through its accreditation by the International Association for Continuing Education and Training (IACET).

For example, Fig. 3 illustrates the potential problem of blowoff if a burner is over-fired. Moving from left to right shows what happens as the air/fuel mixture velocity is increased. The last furnace, on the far right, depicts the danger of going beyond the design firing limit for the burner. The why also better prepares operators to react to new situations that may not have been covered in formal training sessions. The course content included basics that apply to any equipment of this type, along with very detailed and specific information on the equipment in their particular installation. Materials presented include: 1. Combustion and thermal oxidizer basics 2. Safety overview and warnings 3. Overall equipment familiarization

Course design. While the plant is responsible for all safety

practices and training, JZI provided training designed to give operators at the new facility a good idea of both the what and why of operating the TO system and associated equipment. All too often, operators are trained, sometimes hurriedly and haphazardly, by existing experienced operators. The new operators may learn what to do, but not the why behind it. It is also fairly common for long-time operators not to understand some of the basics because they were never taught to them. The why is important because it helps operators better understand the cause and effect that can impact safety, thermal efficiency (and, therefore, operating costs), productivity and pollutant emissions.

FIG. 1

Shintech plant in Plaquemine, Louisiana.

HYDROCARBON PROCESSING NOVEMBER 2010

I 55


SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

4. Detailed walkthrough of equipment 5. Detailed blower, boiler, absorber, scrubber and demister details 6. Detailed walkthrough of P&IDs 7. Drawings review 8. Pre-startup and refractory cure out 9. Normal startup and shutdown 10. Logic demonstration and DCS screens 11. Normal maintenance 12. Troubleshooting 13. Drawings. Each student received a three-ring binder containing the color PowerPoint slides of the course. Adequate room was provided for them to make notes in the manual as desired. Some students received their manuals prior to the start of the class and came prepared with questions to ask. Statistics for all types of training show that retention of the material diminishes fairly quickly after the training has been completed. The student manual can be quickly and easily referenced as often as needed to refresh previously learned information. Although operators do not generally receive their own copy of the operation and maintenance manual, the student manual contains much of the same information, including many of the written operating procedures.

56

FIG. 2

Photo of part of the thermal oxidation system during installation.

FIG. 3

Series of furnaces showing the progression toward blow-off of a burner flame.

I NOVEMBER 2010 HYDROCARBON PROCESSING

Training. The training was conducted over three consecutive days, followed by a fourth day about six weeks later on a couple of specific pieces of the equipment. Although most of the time was spent in the classroom, there were many short sessions spent outside at the equipment to review and emphasize specifics after reviewing the basics in the classroom. The plant had not been started up yet, so the equipment was installed but not operational. While this did not allow the operators to do live training, it did permit operators from all shifts to attend classes together during normal working hours. This produced significant interaction and feedback between participants and with the instructors. Another important aspect of the training was that supervisors were present during most of the sessions, which sent a strong message about the importance of the class. The format of the training was designed to be very interactive. While colorful PowerPoint slides (for example, see Fig. 4) were used to guide the discussion, operators were encouraged to ask questions and make comments at any time. This was encouraged in part through subject-oriented fun games such as poker and bingo. Every time a participant asked or answered a question, they were given a random card from a poker deck. For the poker game, the student with the best poker hand at the end of the day received a prize. For the bingo game, cards were drawn from a deck until someone had enough matching cards to win. The more cards a student had, the more chances of winning, so this encouraged continuous and frequent participation. Other token gifts were also given out during the training as deemed appropriate by the instructors, for example, to a student asking a particularly good question. Short video clips and brief plant visits were used to break up the lecture periods to help keep students engaged in the materials. Videos are particularly powerful when demonstrating potential problems, such as flashback from a burner, that may not have been previously experienced at a particular plant, but which could happen under certain circumstances. This is analogous to airline pilots who train in simulators to react to situations they hope they never encounter, but for which they are prepared to handle just in case. The actual equipment drawings for this plant were used during the training to help familiarize the operators with the equipment and with the operating procedures. To make it even more

FIG. 4

Slide showing the 3Ts of combustion: time, temperature and turbulence.


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SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

realistic, photos were taken of individual components that would pop up on the drawings when clicked. For example, clicking on the symbol of a valve would pop up a picture of the actual valve in the plant as shown in Fig. 5. This personalized the training and made it easier for the operators to connect the drawings to the actual equipment. Results. Identical 15-question pre-tests and post-tests were given to the students to measure learning. The pre-test assessed studentsâ&#x20AC;&#x2122; knowledge prior to taking the class. The average pre-test and posttest scores were 52% and 99%, respectively. The difference between the scores is an indicator of what was learned in the training.

Students were also given a questionnaire at the end of the course to assess their level of satisfaction with the course. Students did not put their names on the forms, although their names were checked off a list to show they completed the evaluation, which is one of the requirements for receiving credits for the course. A five-point Likert scale was used, where 1 = none, 2 = little, 3 = average, 4 = above average, and 5 = great. Students rated each section of the course according to their interest in the topic and its benefit to them. There was also a space to write in any comments they may have had on the topic. Fig. 6 shows the averaged results by interest and benefit for each topic. The results show that, on average, students found all topics to be of above-average interest and benefit. Another part of the questionnaire asked students for written comments on the instructors and material. Some of the instruc5 Interest

BeneďŹ t

6

9

Rating

4 3 2 1 1

FIG. 5

Animated P&ID with a picture of an actual control valve.

FIG. 6

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4

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11

12

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Student ratings of interest and benefit of each course topic.

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PLANT SAFETY AND ENVIRONMENT tor comments included “Very Knowledgeable,” “Excellent” and “Very Thorough.” Some students felt more time should have been spent on startup, shutdown and troubleshooting, and less time on the drawings. Content and coverage are always the challenge with a group of students having a wide range of backgrounds and experiences. Note: All students met the necessary requirements and received CEUs for the class.

1 2

3

SPECIALREPORT

LITERATURE CITED Baukal, C. E., ed., The John Zink Combustion Handbook, CRC Press, Boca Raton, Florida, 2001. Schnelle, K. B. and C. A. Brown, “Thermal Oxidation for VOC Control, Chapter 13,” Air Pollution and Control Technology Handbook, CRC Press, Boca Raton, 2002. Moretti, E. C., “Reduce VOC and HAP Emissions,” Chemical Engineering Progress, Vol. 98, No. 6, pp. 30–40, June 2002.

Outcome. Properly training plant operators is critical to ensure that process equipment is operating safely, while maximizing efficiency and productivity and minimizing pollution emissions. Operators need to understand some basic information about the equipment, as well as the details on their specific installation. Although not always possible, it is particularly beneficial to have all operators together in the same class to enhance discussion and mutual learning. Training should be customized to the needs of the plant and should incorporate techniques such as fun games to promote interaction among the participants and instructors. Ideally, there should be a “hands-on” portion of the training where instructors use the actual equipment during demonstrations. Pretesting and post-testing are effective tools to show that operators have learned the key points in the training. HP

Tim H. Gilder joined Shintech Louisiana, LLC, in 2006 as vinyl chloride monomer

NOTES John Zink Company, LLC (JZC) manufactures thermal oxidation systems used to destroy unwanted wastes.1 The John Zink Institute (JZI) is the training group for JZC and delivers training both at its US headquarters in Tulsa, Oklahoma, and at customer plant sites. JZI works with the plant to determine a suitable course agenda.

Chuck Baukal is the director of the John Zink Institute at John Zink Company,

a b

production superintendent. Prior to joining Shintech, Mr. Gilder served in engineering and supervisory positions at flexible polyurethane foam, furfural, polyvinyl chloride and ethylene dichloride/vinyl chloride monomer production facilities. He earned a bachelor’s degree in chemical engineering from the University of Mississippi.

Dale Campbell, P.E., is a senior design engineer at John Zink Company, LLC, where he serves as the primary resource for incinerator troubleshooting and design in the thermal oxidizer aftermarket group. Since 1973, his primary responsibility has been the detailed design, equipment application, startup, and project management of waste incinerator systems. Mr. Campbell earned a bachelor of science in chemical engineering from the University of Tulsa. Todd Robertson is a combustion service leader at John Zink Company, LLC. He is responsible for thermal oxidizer installation supervision, startup, maintenance, service and training. He retired from the United States Air Force after 23 years of service. Mr. Robertson earned a bachelor’s degree from Embry Riddle Aeronautical University.

LLC in Tulsa, Oklahoma. He has nearly 30 years experience in industrial combustion in a wide range of industries. Dr. Baukal has a PhD in mechanical engineering from the University of Pennsylvania and is a registered professional engineer in the state of Pennsylvania. He has authored/edited eight books on industrial combustion and has 11 US patents.

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Emergency response planning— start at the plant design stage Follow these guidelines for a safer facility R. SAINI, Kuwait National Petroleum Company, Kuwait

A

n optimum level of emergency preparedness can be achieved if emergency response planning is started early and the requirements are incorporated in the plant design. Codes and standards provide basic guidelines but there is also a need to go beyond the books. Emergency response and management requirements should be established and documented, and appropriate provisions should be part of the system. Incidents such as fires and explosions continue to make news every now and then. Such incidents, if not effectively controlled, have serious impact on life, environment, property and business. Adequate and efficient emergency response can minimize the consequences of the incidents, should one occur. Emergency preparedness therefore, continues to be an important element of the overall process safety management system. Fig. 1 shows the components of an emergency preparedness system. It includes the following: • Emergency mitigation and control systems that include process control and emergency shutdown systems, and emergency detection and alarm systems, passive as well as active fire protection systems • Emergency response and management systems that include emergency plans and procedures, communication systems, logistics and resources. Continual reviews, checking, maintenance and testing of emergency preparedness facilities are important to ensure their effectiveness when needed. Overall plant layout and design of the individual facilities, utilities and support systems are normally done based on operational and maintenance requirements. These are done in accordance with the applicable codes, standards and regulations, safety guidelines

and insurance requirements. Emergency planning activity should also be initiated at the project conceptual stage. Fig. 2 demonstrates how the emergency planning team interacts with the plant design team. Emergency preparedness requirements must be established and documented. The following documents should be developed as a part of the project emergency preparedness design package: Emergency scenario register includes: • Emergency scenarios (flammable/toxic liquid/gas leaks, fires etc.) identified for each plant facility based on qualitative risk assessment for each plant facility, considering the products handled and operating parameters and carrying out credibility analysis. • Overall emergency scenario list for the plant along with the recommended emergency response strategies for each type of emergency.

Emergency mitigation and control systems

Assessment and installation

Design stage

FIG. 1

Emergency preparedness

Testing and maintenance

Subsequent periodic adequacy reviews

The emergency scenario register is the basic document for emergency preparedness planning. The data is also used for carrying further quantitative risk assessment and consequence analysis. Fire protection philosophy document should include the design basis for: • Firewater systems—firewater storage and pumping capacity, distribution network design and sizing, fire hydrant design and location • Fire protection requirements for individual facilities such as process areas, storage tanks, operation control rooms, electric substations, cooling towers, waste treatment areas, chemical storage and other buildings • Fire detection and alarm system basic requirements. The philosophy document serves as a basis for detail design as well as for develop-

Emergency response and management systems

Plans and procedures

Training and drills

Emergency response and management teams

Communication system

Equipment and materials

Emergency coordination centers, fire station

Assessment, availability

Testing and maintenance

Emergency preparedness components.

HYDROCARBON PROCESSING NOVEMBER 2010

I 61


PLANT SAFETY AND ENVIRONMENT

Codes, standards, regulations

ing a fire protection manual for subsequent use. The fire protection manual includes details of the actual as-built fire protection systems installed, along with testing and maintenance requirements. Pre-incident planning for a particular plant facility/area includes: • Listing all types of potential emergency incidents and possible escalations • Establishing an incident control strategy and resource requirement, checking the adequacy of the fixed emergency control systems to control the incident • Assessing external emergency response requirements—manpower, equipment and materials as necessary to supplement the fixed systems and also as an alternate in case of fixed-system failure • Developing an emergency response equipment deployment plan, including fire hose laying requirements according to the availability of fire hydrants. The pre-incident plans should be developed for all plant facilities/areas. The worstcase scenario requirements for each type of emergency are used to verify the design of the common emergency response facilities. For example, pre-fire plans should be used to review and update, if necessary, the fire protection design philosophy and adequacy of the fire protection and emergency response systems, including the following: • Firewater system design—design flow and pressure, water storage, distribution network layout and sizing, to meet the worst-case fire scenario demand • Review and update of the number and location of fire hydrants • Number and types of emergencyresponse vehicles and materials.

FIG. 2

62

The information from pre-incident plans is also used for: • Identifying gaps and possible improvements • Checking the adequacy of the plant drainage system to handle firewater effluent as well as any hazardous material discharge. Emergency response and management planning should also be started at an early stage and the location and design requirements for the following should be established: • Emergency communication systems • A site emergency response center for identified incidents such as fires, oil spills, hazmat violations, chemical spills and personnel injuries • A holding area for external response teams • Evacuation requirements • An emergency management center— emergency management team work station. Emergency communication requirements include: • Automatic detection and alarm systems, manual call points • An emergency communication control and emergency response team dispatcher center • Personnel call-out communication • Public address and siren systems. Emergency evacuation requirements: • Local assembly points—pre-identified areas near each occupied building or work area where the occupants gather in case of an emergency. They wait for headcount and further instructions. • Main gathering areas and shelter-inplace—pre-identified area in a safe loca-

Perimeter fencing, facilities arrangement, safety distances

Overall plant layout

Emergency response access

Life safety, passive fire protection, etc.

Facility design

Evacuation, assembly points/shelters

Communication systems Plant drainage

Emergency coordination centers

Other systems

Firewater systems

Emergency communication

Active fire protection systems – philosophy/ manual

Mobile fire-fighting equipment

Process control systems

Hydrant design, spacing

Emergency resources (worst-case scenario)

Emergency response plans

Emergency scenario identification

Incident control strategies

Pre-fire plans

Emergency planning at the design stage improves emergency preparedness.

I NOVEMBER 2010 HYDROCARBON PROCESSING

Emergency planning

SPECIALREPORT

tion where the evacuees might stay for a longer duration in case of a major plant emergency. Such areas should provide protection against severe weather conditions and be equipped with basic facilities such as drinking water, toilet, first-aid and communication. Large cafeteria dining halls, auditoriums, basements, etc., if available, may be considered for use as gathering areas during emergency evacuations. Also, consideration should be given to designing buildings as a shelter-in-place for use in the event of a toxic gas release emergency. It is advisable to identify additional areas to be used as alternate gathering areas. • Evacuation routes—leading to local assembly points and further to main gathering areas to be identified and displayed prominently at appropriate locations, along with necessary instructions and guidelines. Some important areas requiring special consideration in the plant layout, fire water system design, pre-fire planning and emergency response planning are: • Large hydrocarbon storage tanks • Wind direction • Plant drainage system. Hydrocarbon storage tank fires have always been a challenge for firefighters. Tank fires have potential to escalate, if not controlled in the initial stages, and the consequences may be disastrous. This is more so with large storage tanks and tanks containing light products and products with ‘boil over’ potential, e.g. crude oil. Special consideration should be given while developing tank farm layout, especially when large tanks are part of the design. Roads should be provided on at least three sides of large-diameter tanks so that sufficient access is available for fire fighting appliances. Maximum cumulative volume of the flammable liquid stored in all the tanks in the dyke should be limited to reduce risk. The Institute of Petroleum (IP) Safety Code of Practice, Vol. 3 provides guidelines in this regard. Dykes or bunds are important aspects of designing tank farms. Codes permit dykes made of earthen, concrete or masonry construction. Concrete or masonry dykes help save space and require low maintenance as compared to earthen dykes. However, their use should be limited to low heights (roughly 1 m) wherever possible. Concrete dykes restrict the ventilation and dead-air pockets that are formed close to the dyke base where flammable vapors can accumulate. This is also a hindrance to firefighters and it may be difficult for a person to escape during an emergency. The dyke


PLANT SAFETY AND ENVIRONMENT layout dimensions can be worked out by volume calculations, code requirements and accesses. Wherever possible, the following points should be considered: • Maximize the tank spacing within the same dyke. However, the tank spacing with respect to the adjacent dykes should not be ignored. • Space between the dyke and the tank shell should be such that the liquid jet from any potential leak from a hole in the tank shell is contained within the dyke. • Reach of the foam thrown from portable foam equipment is an important consideration. Therefore, the road to tank distance should be optimized, without compromising other safety requirements, to ensure effectiveness of the portable foam equipment to be deployed as per the plan. Pre-fire plans for storage tanks should cover many scenarios including small leaks, rim seal fires and the possible escalation to full open-top fire. A fixed-foam application and cooling systems attached to the tank are the most effective means for controlling tank fires. Floating roof tanks are normally provided with fixed foam pourer systems for rim seal fire protection only. However, full-surface open-top tank fires, though not very common, can and do occur if the floating roof sinks. In absence of the fixed foam systems controlling full surface open top fires in the large floating-roof tanks will require foam application by using portable and mobile equipment. Using portable and mobile equipment may also be necessary when the fixed systems, where provided, might get damaged or become inoperative. Therefore, it is necessary to ensure that adequate provisions are made in the fire protection systems design. For calculating the overall design firewater demand, water requirements for a foam application over a full open-top surface area, tank shell cooling for the tank on fire and exposure protection of the adjacent tanks should be considered. It must be noted that the foam application rate is considerably higher when foam is applied using portable means as compared to the foam application using fixed systems. Recent research by specialist fire-fighting agencies has recommended higher application rates than those specified in the codes. High discharge foam monitors with large-diameter hoses should be used for tank fire fighting. Deploying a large number of the smaller-capacity foam monitors with conventional 65-mm fire hoses may not be practical. Also, a foam monitor deployment

plan needs to be prepared, considering the prevailing wind direction. This implies that one side of the tank (downwind) will not be available for placing foam equipment. The wind direction changing during fire fighting is still another constraint. Alternate locations must be considered. An adequate number of fire hydrants should be available at strategic locations to feed the foam equipment. Fire hydrants should have outlets to enable connecting of large-diameter hoses for the purpose. Thus, the pre-fire

SPECIALREPORT

plan should include the estimated number of fire hoses along with the foam equipment requirements. Sufficient quantity of foam concentrate should be made available. Wind direction is another important factor that requires proper attention during plant design and emergency response planning. Drifting and pattern of a hazardous vapor cloud depends on the wind direction. In case of a toxic gas release the occupied areas downwind of the leak source will be affected. In case of a flammable gas

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SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

leak, there are chances of the drifting flammable vapor cloud coming in contact with a source of ignition on the downwind side, with potential danger of an explosion and fire. A plant layout design team must consider the prevailing wind direction, while locating the potential critical sources of the hazardous gaseous leak, with respect to the populated areas/sources of ignition inside as well as outside the plant boundaries. It is important to understand the prevailing wind direction concept. The aim is to locate the potential toxic/flammable gas leak sources in such a manner that the chances of the vapor/gas cloud drifting toward the ignition source or the populated area are minimal. Some sites have a clear single prevailing wind direction throughout the year, while other sites may have different prevailing wind directions during other times of the year. Therefore, designers must study the wind rose diagrams carefully and arrive at a direction that has the minimum chance on wind blowing in the wrong direction. Such direction is often termed as the cross-wind direction. Drains and sewers also have an impact on plant fire safety. Industry has witnessed the consequences of the plan drainage sys-

tem designed without proper consideration to fire safety, such as: • Fire spreading through underground oily sewers • Explosion in sewers • Fire spreading due to an oil layer on the open drains • Drains overflowing and water logging during fire fighting. Fire safety considerations related to plant drainage systems are: • Drainage systems must be designed to handle the firewater discharge expected during fire emergencies • Oily sewers should incorporate liquid seals to stop propagations of fire through vapor spaces in the sewer • Open drains, designed for non-oily discharges in normal operations still have chances of carrying oily effluent in abnormal situations, and should also be provided with liquid seals at certain intervals/locations. Conclusion. Effective emergency response is essential to minimize the consequences of an incident in the petroleum industry. Plant design must incorporate the emergency response requirement, in addition to the normal plant safety requirements

laid out in the code and standard guidelines. Practical aspects of emergency planning and response should be incorporated in the plant design, along with consulting emergency response specialists. HP Ramesh Chand Saini is a Health Safety and Environment professional with over 35 years of experience in emergency preparedness and response systems in the hydrocarbon processing industry. He currently works at Kuwait National Petroleum Company (KNPC) and his responsibilities include developing and updating plans and procedures for fire safety engineering and management, emergency response and crisis management systems; and also ensuring fire safety adequacy of the existing plants as well as the new projects. Mr. Saini has initiated a number of projects for upgrading fire safety facilities in refineries. After graduating from the Indian Institute of Technology (IIT) in Kanpur, India, he worked for 13 years for Engineers India Limited (EIL) in its project engineering team in the area of safety and fire protection for refineries and petrochemical projects. Mr. Saini is a member of the Society of Fire Protection Engineers (SFPE), a certified HAZOP leader and ISO 9001 lead auditor and has participated in a number of HAZOP studies and audits of fire safety facilities at oil installations across Kuwait. He is a member of a number of fire incident investigation teams. Previously, Mr. Saini was a member of the fire committee of the Bureau of Indian Standards, involved in developing a number of Indian standards on fire safety. He has also been briefly associated with the Oil Industry Safety Directorate of India for developing safety standards.

The Fundamentals of Piping Design By Peter Smith 262 pages • Hardcover • Pub date: April 2007 ISBN: 978-1-933762-043 • Price: $175 Written for the piping engineer and designer in the field, this first part of the two-part series helps to fill a void in piping literature, since the Rip Weaver books of the ‘90s were taken out of print.

Advanced Piping Design By Rutger Botermans and Peter Smith 250 pages • Hardcover • Pub date: May 2008 ISBN: 978-1-933762-18-0 • Price: $175 An intermediate-level handbook covering guidelines and procedures on process plants and interconnecting piping systems.

The Planning Guide to Piping Design By Richard Beale, Paul Bowers and Peter Smith 300 pages • Hardcover • Pub date: September 2010 ISBN: 978-1-933762-37-1 • Price: $175 The Planning Guide to Piping Design covers the entire process of planning a plant model project from conceptual to mechanical completion, and explains where the piping lead falls in the process along with his roles and responsibilities.

To place an order, visit www.gulfpub.com or call +1 (713) 520-4426. Select 165 at www.HydrocarbonProcessing.com/RS 64

Select 166 at www.HydrocarbonProcessing.com/RS


PLANT SAFETY AND ENVIRONMENT

SPECIALREPORT

Optimized fired heater control Residual oxygen measurement principle lowers emissions and improves efficiency A. J. MOURIS, Hobré Instruments, Purmerend, The Netherlands

T

ighter emission regulations and high energy costs pose new challenges to control systems for fired heaters. Rapid changes in fuel gas heating value, air demand and composition are typical for applications in oil refineries, chemical plants and many other sites. Traditional feedback control based on temperature, stack oxygen and combustibles measurement is not quick enough to handle rapid changes effectively. This short coming is typically addressed by controlling the excess air set point with a certain safety margin. Unfortunately, this approach prevents the emission of unburned components while increasing CO2 emission due to poor fuel economy—air is heated unnecessarily and heat transfer efficiency is reduced. NOx formation is promoted as a result of higher oxygen levels in the combustion process. For these reasons, feed forward control of the air/fuel ratio is gaining more attention. Properly selecting and installing the fuel gas property analyzer and using the right control parameters are essential to get the best results.

Control parameters. The control system philosophy of fired heaters varies depending on the requirements and heater or boiler design. However, in all cases, the furnace’s thermal load and the air/fuel ratio are two critical parameters that must be monitored and controlled. Heat load control. Depending on the

control system design, the Wobbe Index (WI), the heating value and gas density may be required as input(s). The heating value is the amount of heat produced when a unit volume or fuel mass is burned stoichiometrically. The higher heating value includes the heat of water condensation formed in the combustion process; the lower heating value does not. The WI (defined as the

heating value of a gas divided by the square root of its specific gravity) is a measure of the interchangeability of fuel gases when introduced into a heater via a burner with a fixed differential pressure. Two gases with the same WI will deliver the same amount of heat into a combustion process per unit of time, regardless of the composition. To clarify this concept, consider the following fuel gas cases: • Case 1—40% methane and 60% hydrogen (by volume) • Case 2—58% methane and 42% nitrogen. The lower heating value by volume of these two gases is the same, i.e., 20.82 MJ/Nm3. The WI however, is 40.55 MJ/ Nm3 for Case 1 and 24.39 MJ/Nm3 for Case 2! This means that the amount of heat delivered per unit of time through the same burner will be 40% lower in the second case.

taken into account. If large fluctuations in the fuel gas composition are expected, the signal from a WI analyzer or calorimeter is used for correcting the air/fuel flow ratio. Typically, the assumption is that there is a proportional relationship between heating value and air demand. Whereas this is correct for hydrocarbon-based fuel gases like natural gas, for fuel gases containing significant percentages of hydrogen, olefins, CO2 and/or oxygen this approach fails. The following are fuel gas cases: • Case 1—100% hydrogen • Case 2—88.5% methane and 11.5% nitrogen. The WI for both gases is the same, i.e., 40.9 MJ/Nm3. The Combustion Air Requirement Index (CARI) is defined as the stoichiometric air demand divided by the square root of the relative gas density. Results: 9.0 for Case 1 and 10.9 for Case 2. This means that if the fuel gas composition changes from hydrogen poor to hydrogen-rich composition the excess air may be controlled 20% too high. Please note that, instead of WI and CARI, a similar case can be construed for heating value and air demand; this follows from the definitions:

Air/fuel ratio control. The combus-

tion air flow supplied to an industrial furnace is typically linked to the fuel gas flow. In smaller installations, this may be a mechanical link; in larger installations, air and fuel gas temperature and pressure are

Flowmeter Sample

SG

Pressure reducer

Low

SG

Bypass Bypass Vent

High Manifold ZRO2 cell Restriction Heat exchanger

PI

Air station FIG. 1

Mixing chamber

Measuring oven

Drain

Booster

Typical residual oxygen content analyzer schematic.

HYDROCARBON PROCESSING NOVEMBER 2010

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Select 60 at www.HydrocarbonProcessing.com/RS


PLANT SAFETY AND ENVIRONMENT WI = Heating value/√(Specific gravity) CARI = Air demand/√(Specific gravity) Residual oxygen content analysis. In a typical residual oxygen content

analyzer sample, gas is continuously mixed with combustion air under controlled conditions followed by catalytic combustion in an electrically heated furnace. The residual oxygen content in the flue gas is measured with an accurate and reliable zirconium oxide sensor. In the control unit, the following combustion parameters are calculated from the oxygen signal and the (optional) density signal: WI, CARI, calorific value (or BTU) and specific gravity. The concept was first explored in the US, but the European gas distribution companies Gaz de France and Dutch Gasunie have really optimized the benefits. Their prime objective was to develop an instrument that was as fast as possible for optimizing natural gas blending operations to meet grid entry specifications. However, in the last decade the technology has also proven to be very suitable for fuel gas, vent gas, flare gas, biogas and steel plant offgas applications. Instrument installation and selection. When the decision is to install an

analyzer for measuring the heating value and/or WI for feedforward fuel and air/fuel ratio control, the following requirements should be fulfilled: • The analyzer should be as fast as possible. It doesn’t make sense to install a calorimeter with a 20-sec response time if changes occur within seconds. With residual oxygen technology, a response time of less than 5 sec is achievable. • Signal noise should be as low as possible. High signal noise levels require smoothing of the signal, typically by averaging. As a consequence, the response from the control system to a step-change will be slower. A residual oxygen content analyzer offers a repeatability of 0.05% of measured value. • Local installation should be close to the sample tap point. Ideally, the fuel gas heating value and air demand signal should be available before the fuel gas leaves the burner tip. This means that the travelling time of the fuel gas from the sample tap point to the burner should be longer than the traveling time from the sample tap point to the analyzer plus the analyzer response time. Outdoor installation in hazardous areas is not a must, but it often follows from the requirment of installation close to the sample tap point.

• The sample handling system should have minimal internal volume. Although a fast-loop theoretically can compensate for any dead volume in the system, this will result in excessive venting and/or flaring of fuel gas. Ideally, the analyzer has an integrated sample conditioning system and requires no additional external sample handling. • A combustion air requirement signal should be available. As discussed previously, the heating value or WI can be

SPECIALREPORT

poor indicators for the air demand in fuel gas applications. Therefore, it is important to select an analyzer based on the residual oxygen content principle that stores separate calibration lines for WI/heating value and CARI/air demand. • The analyzer’s rangeability should match all possible cases. Typically, the analyzer should be able to handle large fluctuations in the fuel gas composition. Residual oxygen content analyzers analyze fuel gases of all possible compositions in the 0–120

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HYDROCARBON PROCESSING NOVEMBER 2010

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SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

■ In a typical residual

oxygen content analyzer sample, gas is continuously mixed with combustion air under controlled conditions followed by catalytic combustion in an electrically heated furnace.

MJ/Nm3 (0–3,000 BTU/SCF) range without the risk for flame-out or overheating. • Thorough application review. Each application is different and proper review is essential. Issues that should be considered include: º Calibration gas selection—the “right” calibrated gases give the best accuracy in most cases, do not contain many components (not more than two preferably) and allow sufficient filling pressure even when ambient temperature may be low.

º Parameters measured—besides CARI and WI, specific gravity, heating value and air demand may be required. º Fuel gas hydrocarbon and/or water dew point—it is not uncommon that fuel gas is taken from a knock-out vessel. Care must be taken that condensation does not take place in sample lines or inside the analyzer. Ideally, all sample wetted parts should be heated 10°C–20°C above the maximum expected dew point. º Sulfur content and presence of other corrosive components— wrong material selection can rapidly corrode and clog an analyzer. Proper component selection and analyzer design enables continuous operation even when more than 10% sulfur is present. º Overall response time—a lag time analysis from the sample probe tip to the analyzer signal output should be provided to ensure compliance with the requirements. This is especially important when high pressure gas lines must be analyzed. º Ambient temperature range and hazardous-area certification requirements—these must be considered. Conclusion. Feedforward control of

fired heaters utilizing fuel gas with variations in composition and properties may result in considerable improved combustion efficiency. Emissions may be reduced and product quality and equipment life time can be improved. Online analyzers based on the residual oxygen content method have the potential to deliver the parameters to be measured within seconds. However, a successful implementation of the technology requires a clear understanding of what is to be measured and how the signals are used in the control loop. Also, the location of installation and a properly designed sample handling system are key factors for success. HP

MODERNIZING Level Measurement

Sump Level Solution Sump level measurements and pump control are often a maintenance

1

challenge. The VEGAPULS through-air

2

radar and VEGAFLEX guided wave

3

radar provide reliable continuous level measurements. The VEGAMET 391 supplies up to six relays, offering

LITERATURE CITED API RP 556, First Edition, May 1997. Driedger, P., “Controlling fired heaters,” Hydrocarbon Processing, April 1997. Physical Properties of Natural Gases, N.V. Nederlandse Gasunie, 1980. s

80-90%

10-20%

lead/lag pump control and giving power to the measuring instrument.

www.ohmartvega.com info@ohmartvega.com 800.FOR.LEVEL

Select 168 at www.HydrocarbonProcessing.com/RS 68

Albert Mouris is technical director with Hobré Instruments BV. He has more than 15 years of experience in process analysis and sampling systems. Previously, Mr. Mouris worked as an energy market analyst with the Dutch gas incumbent Gasunie. He received a chemical engineering degree from the University of Twente.


PLANT SAFETY AND ENVIRONMENT

SPECIALREPORT

Consider real-gas modeling for turboexpanders New visualization methods expose problems with traditional designs K. KAUPERT, OC Turboexpanders, Irvine, California

E

very year turboexpanders generate millions of Euros in revenue for hydrocarbon processing plants by removing heat from gas streams, also known as the “turboexpander refrigeration benefit.” To maximize this financial benefit, accurate gas dynamic performance predictions for turboexpanders are a necessity. This requires an accurate thermodynamic equation of state that uses a real-gas model. But, which real-gas model is best? For example, an ideal gas assumption can cause horribly wrong performance predictions due to gas compressibility at high pressures and low temperatures.1 As a result, all turboexpander manufacturers use real-gas models in their simple gas dynamic sizing predictions. However, beyond the simple sizing predictions, advanced turboexpander manufacturers continue to apply real-gas modeling in the detailed computational fluid dynamics (CFD) design of its turboexpanders. This permits turboexpanders, such as shown in Fig. 1 to attain high efficiency levels. Historically, the application of real-gas models in commercially available CFD packages has been problematic or even non-existent. This is due to: 1. Most turbomachinery manufacturers (e.g., those producing gas turbines or turbochargers) are content with an ideal gas law or a simple real-gas model since their compressibility effect is modest and the flow is single phase. 2. Increased computational time is required when a real-gas model is applied in CFD (slower code). 3. Numerical robustness is decreased when a real-gas model is applied in CFD (code can more easily diverge). 4. Difficulty modeling two-phase flow in a wet gas expansion region. The turboexpander manufacturer is, therefore, confronted with somewhat unique challenges, as the expander inlet gas can exhibit substantial compressibility while the expander outlet gas can exhibit two-phase wet gas flow.2 For example, Fig. 2 shows an expander impeller connected to a rotating assembly. This particular impeller is subjected to both compressible gas and wet gas, which requires accurate real-gas modeling in CFD to maximize expander efficiency. But there are many real gas models available in the open literature. So, the question again arises, which model is best? In this article, we will assess several real-gas models used in CFD for turboexpanders.

Ideal gases and liquids. The simplest, most popular equation of state for gases is the Ideal gas law; it states: P = ␳RT (1) For liquids, the bulk modulus and coefficient of thermal expansion combine for the equation of state given as:

FIG. 1

Turboexpander designed with real-gas modeling in CFD.

FIG. 2

Expander side of a turboexpander rotating assembly.

CONSIDERING THE EQUATION OF STATE

An equation of state relates a fluid’s state variables.3 For turboexpanders, the equation of state is a thermodynamic equation, which mathematically describes the interaction of the macroscopically measurable properties of the process fluid (e.g., the thermodynamic variables of pressure, temperature, density and composition).

HYDROCARBON PROCESSING NOVEMBER 2010

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SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

Kl = dP/(d␳/␳) and ␣l =(1/V)dV/dT (2) Over limited ranges of pressure and temperature and without phase changes, these two equations ( Eqs. 1 and 2) give reasonable property predictions (e.g., the ideal gas law for ethane at less than 5 bar pressure and higher than 100°C temperature gives less than 2% error on the density). But turboexpanders routinely handle gases outside the limited range of the Ideal gas law and with two-phase flow. In essence, real gas and liquid modeling is obligatory. Real gases and liquids. The need for accurate equations of

state has resulted in an abundance of real-gas models in the open literature. So numerous are the real-gas models that they could not possibly all be evaluated in this short article. However, these five real gas equations of state do find widespread application: 1. Peng-Robinson (PR) 2. Soave-Redlich-Kwong (SRK) 3. Benedict-Webb-Ruben-Starling (BWRS) 4. Lee-Kesler-Plöcker (LKP) 5. GERG. These five equations of state are compared here with CFD results through the expander side of a turboexpander. The math-

ematics of these equations of state are presented elsewhere; they are large in size and require detailed explanation, among them Refs. 4 and 5. The American Gas Association (AGA-8) is not evaluated here as it poorly predicts the weight fraction of liquid in the wet-gas region at the expander outlet. Another very common and simple real gas equation of state utilizes the compressibility factor, Z and is called the universal gas law, given as: P = Z␳RT (3) Unfortunately, the determination of Z for gas mixtures in turboexpanders is very difficult; this equation (containing Z ) is not evaluated here. CFD WITH REAL-GAS EQUATIONS OF STATE

CFD using a real-gas equation of state is a time-consuming task. In the past, this author has pursued three methods to incorporate real-gas equations in CFD. The first method is to program the real-gas equation of state and patch it into the CFD code. While this method is the most straightforward, it frequently leads to numerical convergence difficulties in the transition region from single-phase gas to two-phase wet gas. This concern can be addressed with numerical damping routines that are ingeniously applied by some CFD vendors. The second method is to use the real-gas equation of state to prepare a set of “look-up” tables in the desired pressure and temperature range of the gas. This method is more tedious for pre-processing the CFD but tends to avoid convergence problems while running the code. A third method to follow uses the tables from the second method to create a set of polynomials to approximate the equation of state. But the third method can lead to large errors outside a prescribed temperature and pressure range and convergence difficulties for the CFD. TABLE 1. Gas conditions for the expander side of the NGL hydrocarbon example Gas dynamics Job: Example NGL fractionation Component

FIG. 3

Expander, mol%

16.0430

81.000%

Ethane

30.0700

11.000%

Propane

44.0970

5.000%

i-Butane

58.1230

0.800%

n-Butane

58.1230

0.900%

i-Pentane

72.1500

0.344%

n-Pentane

72.1500

0.100%

255

n-Hexane

86.1770

0.050%

245

C7+

110.00

0.002%

Nitrogen

28.0134

0.800%

235

Carbon dioxide

44.0100

0.003%

225 K

Water

18.0153

0.001%

Density as a function of pressure and temperature from a real-gas model and used in a “look-up” table for CFD.6

275 K 5.5 7.0 MPa 6.0 5.0 4.5 4.0 6.5 3.5 MPa

Enthalpy

MW

Methane

a b

Total Given process conditions Molecular weight

d c

Entropy

FIG. 4

70

Enthalpy vs. entropy diagram for a hydrocarbon gas example.

I NOVEMBER 2010 HYDROCARBON PROCESSING

100% Rated case expander 20.087

Inlet pressure, P1, MPa a

6.91

Inlet temperature, T1, °C

0.2

Outlet pressure, P2, MPa a Mass flow, kg/sec Volume flow, Nm3/h

3.48 13 52036


PLANT SAFETY AND ENVIRONMENT Most CFD vendors have followed the direction of the second method or at least opened up their code for input from such “lookup” tables. It is the second method that the author has applied here. Essentially, the real-gas equation of state is used to generate the thermodynamic “look up” tables for P=P (u,␳), T =T(u,␳), P = P (h,s), ␳=␳ (h,s), s=s(h,P), h=h(s,P), u=u (P,T) and ␳=␳ (P,T) where the variables are: P Pressure T Temperature ␳ Density u Internal energy h Enthalpy s Entropy. An example of such a “look-up” table is plotted in Fig. 3, taken from Ref. 6.

SPECIALREPORT

surements for similar gas compositions and conditions, the GERG model was found to be most accurate for this gas case. Although, the LKP and BWRS models would be acceptable as well. Results from applying the GERG model through the use of generated property tables in the CFD are seen in the values given in Figs. 5 and 6 at the expander design point. Fig. 5 shows the distribution of temperature as predicted with the GERG real-gas equation of state in the natural gas mixture. A rapid temperature change is seen through the expander nozzles; it is typical for turbo expanders, as the temperature decreases due to the acceleration of the flow. In Fig. 6, the relative velocity vectors in the expander impeller are seen, again as computed with the GERG real-gas model. The overall image is shown on the left side of the figure and a zoom at the expander impeller trailing edge is seen at the

Example: CFD for real gases in natural gas liquids.

To protect client confidentiality, a generic gas composition and condition were selected for this example. However, both gas composition and condition are representative of a natural gas liquids feedstock to the expander side of a turboexpander. The gas composition and condition are seen in Table 1. The enthalpy vs. entropy diagram is seen in Fig. 4 along three lines for the gas expansion. The red line “a-b” represents an isenthalpic gas expansion as would be experienced through a Joule-Thompson valve. The blue line “a-c” represents an isentropic, or perfect, expansion of the gas. The green line “a-d” represents the gas expansion through the turboexpander and indicates a lower outlet temperature than the isenthalpic red line expansion. This highlights the benefit of using a turboexpander vs. a JouleThompson valve for the gas expansion—a colder gas outlet temperature with the green line and heat removal from the expander gas stream. Table 2 shows the outlet gas conditions at the point “d” FIG. 5 CFD results for the temperature reduction in the expander compared between the ideal gas and five real gas models. It is seen per GERG real-gas model. immediately that the ideal gas model gives the lowest expander outlet temperature along with the largest expander wheel (impeller) output power. However, the compressibility Z is seen as 1.0, which is not realistic for the expander inlet or outlet gas conditions. Accordingly, applying an ideal gas model would lead to a substantial modeling error for this gas. Among the five real-gas models, the PR model gives the lowest expander outlet temperature, lowest output power, and lowest efficiency. The GERG, LKP and BWRS FIG. 6 Left: CFD results using the GERG real-gas model for the relative velocity vectors in models all give similar outlet temperature, the expander impeller (wheel). Three blades are shown along the shroud showing the output power and efficiency, as well as, simiparallel flow to the blade at the outlet. A zoom is shown to the right at the impeller lar weight liquid percentages. Based on meaoutlet. TABLE 2. Comparison table of CFD results from five real gas models at the expander outlet Computed outlet conditions Outlet temperature, T2, °C Zin/Zout Specific enthalpy ⌬Hs, kJ/kg Expander isentiopic efficiency, % Expander impeller power, kW Outlet weight liquid, %

Ideal-gas model Ideal

Real-gas model GERG

Real-gas model PR

Real-gas model LKP

Real-gas model SRK

Real-gas model BWRS

–33.7

–28.8

–29.4

–28.9

–28.3

–28.7

1.0/1.0

0.72/0.78

0.69/0.76

0.71/0.79

0.70/0.77

0.71/0.78

86.84

54.03

52.12

54.37

52.52

53.69

82.7

85.3

84.6

85.4

84.8

85.2

933.2

599.1

573.2

603.6

579.0

594.7

unable

13.60

12.74

13.52

13.32

13.64

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SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

TABLE 3. Generalized recommendations regarding applicable equations of state Gas type

GERG

PR

LKP

SRK

BWRS

Hydrocarbon

+/â&#x20AC;&#x201C; 2%

+/â&#x20AC;&#x201C; 5%

+/â&#x20AC;&#x201C; 2%

+/â&#x20AC;&#x201C; 3%

+/â&#x20AC;&#x201C; 2%

Air

+/â&#x20AC;&#x201C; 2%

+/â&#x20AC;&#x201C; 5%

+/â&#x20AC;&#x201C; 3%

+/â&#x20AC;&#x201C; 3%

+/â&#x20AC;&#x201C; 2%

CO2

+/â&#x20AC;&#x201C; 3%

+/â&#x20AC;&#x201C; 6%

+/â&#x20AC;&#x201C; 4%

+/â&#x20AC;&#x201C; 4%

+/â&#x20AC;&#x201C; 1%

H2

+/â&#x20AC;&#x201C; 4%

+/â&#x20AC;&#x201C; 7%

+/â&#x20AC;&#x201C; 4%

+/â&#x20AC;&#x201C; 2%

+/â&#x20AC;&#x201C; 5%

not good

not good

not good

not good

not good

NH3

Note: the PR equation of state was the least accurate for all gas types but it is also the simplest to apply. For ammonia, none of the five real-gas models performed satisfactorily and in-house models are still relied upon.

right, showing that the relative velocity vectors closely follow the blade surface without any recirculation zones. Such visualizations are the key to optimizing the gas path and efficiency of expanders by reducing any unwanted entropy generation. Real-world view. Real-gas modeling is needed for accurate gas dynamic performance predictions in turboexpander CFD to optimize expander efficiency. As the five models have shown, different results are attained by applying different real-gas models. It is important for turboexpander manufacturers to use test results and to also obtain detailed field feedback. Together, this will allow selecting the real-gas model best suited to a particular application. In this article, only one example was presented. It is not wise to generalize on the basis of just this one example. The scope is widened with other published or in-house data. Table 3 offers a number of generalized recommendations as to which equation of state should be considered for CFD in modern turboexpanders.

Finally, interested readers are always encouraged to review comprehensive texts on turbomachines, (such as Ref. 7) or more elementary books on turboexpanders (Ref. 8). As energy conservation has become one of the worldâ&#x20AC;&#x2122;s foremost priorities, the importance of efficient turboexpanders continues to increase. HP 1

2 3 4 5

6 7 8

LITERATURE CITED Beinecke, D. and K. LĂźdtke, Die Auslegung von Turboverdichtern unter BerĂźcksichtigung des realen Gasverhaltens, VDI-Berichte 487, VDI-Verlag DĂźsseldorf, pp. 271â&#x20AC;&#x201C;279, (in German), 1983. Kaupert, K. A., â&#x20AC;&#x153;Design of Two-Phase Flow Air Separation Turboexpanders,â&#x20AC;? Cryogenic Technology Journal China, Vol. 1, pp. 47â&#x20AC;&#x201C;52 (in Chinese), 2010. Zemansky, M. W., Heat and Thermodynamics: An Intermediate Textbook, McGraw-Hill Book Co., 6th edition, 1981. Modisette, J. L., â&#x20AC;&#x153;Equation of State Tutorial,â&#x20AC;? Pipeline Simulation Group (PSIG), Paper 0008-2000, 2000. Kunz, O., R. Klimeck, W. Wagner and M. Jaeschke, The GERG-2004 WideRange Equation of State for Natural Gases and Other Mixtures, VDIBerichte 557 Reihe 6, VDI-Verlag DĂźsseldorf, 2007. Numeca, Numeca Fine Users Manual Version 6.1-1, Numeca International, February 2003. LĂźdtke, K., Process Centrifugal Compressors, Springer Verlag, Heidelberg, 2004. Bloch, H. P. and C. Soares, Turboexpanders and Process Applications, Gulf Publishing Co., Houston, ISBN 0-88415-509-9, 2001.

Dr. Kevin Kaupert is the director of technology at OC Turboexpanders. He holds a doctorate in turbomachinery engineering from the ETH Zurich Swiss Federal Institute of Technology. He has over 25 years of experience in turbomachinery for cryogenics, power generation and aerospace applications.

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PLANT SAFETY AND ENVIRONMENT

SPECIALREPORT

Predictive emissions monitoring helps reduce stack air emissions New technology reduces compliance costs while optimizing operations R. HOVAN, Rockwell Software Environmental Solutions, Rockwell Automation, Austin, Texas

T

he oil and gas industry is under unprecedented environmental scrutiny. Already struggling to cope with increasingly rigorous air quality regulations, the industry’s been battered in recent headlines about oil spills from the Gulf of Mexico to Michigan’s Kalamazoo River. To protect both the environment and company reputations, producers must do far more than prevent potential environmental crises. They need to employ solutions and invest in technology to make oil and gas production cleaner. Today, oil and gas processers across the globe are facing stringent emissions monitoring, reporting and reduction goals from national regulatory bodies, as well as an ever-growing patchwork of state and regional regulations. Carbon dioxide (CO2)—the primary greenhouse gas (GHG)— and nitrous oxide (N2O)—a GHG largely blamed for acid rain and ozone depletion—are among the top targets for reduction. Traditional emissions monitoring and management relies heavily on hardware, manpower and a variety of reporting forms and spreadsheets. Today’s and tomorrow’s regulations require a level of accuracy, timeliness, consistency and security that legacyemissions monitoring systems cannot easily provide. The good news is that cost-effective, software-based predictive-emissions monitoring systems (PEMS) and environmental management applications already are available to help reduce emissions. These systems use existing plant process-monitoring equipment to measure emissions factors in real time, predict future emissions scenarios and compile emissions reports, thus helping to reduce the cost of regulatory compliance. Any business manager who’s well-versed in efficient production will recognize the potential uses for such accurate, immediate data collection and analysis. Oil and gas producers already facing increasing capital and operational costs can use this data for more than achieving regulatory compliance. It can also help reduce energy usage and improve productivity and business performance. The technology for energy efficiency that actually improves competiveness is here and it’s proven. Regulatory ramp-up. Petrochemical and natural gas plants face new emissions-related regulations that require immediate action. New rules aim to prepare the industry for emissions reduction by focusing on monitoring and reporting improvements. Plants that produce combined emissions (from combustion, flares, fugitive emissions and vents) equal to or greater than 25,000 metric tons of CO2equivalent (CO2e) annually must monitor emis-

sions and provide comprehensive reporting on an annual basis. This is according to the 2009 GHG Mandatory Reporting Rule from the US Environmental Protection Agency (EPA). Simply put, the new regulations require that producers know and report precisely what is coming out of their smokestacks. The rule requires monitoring systems to be in place by Jan. 1, 2011. The first annual emissions report is due March 31, 2012. After enacting the 2009 rule, the EPA recognized that implementing a traditional monitoring system might be cost-inhibitive for some producers. As a result, the EPA adjusted a number of the requirements outlined in the initial rule to help reduce the burden. Under the updated rule, facilities may measure emissions through engineering estimates and emission-modeling software. Despite these changes, oil and gas producers have encountered many challenges in trying to comply with requirements. Processing is a complex operation and emissions can vary widely, depending on a spectrum of variables ranging from the composition of the fuel feedstock to weather conditions. Many producers today are finding that they don’t know if they’ve surpassed the regulated limit for specific emissions until they prepare their required reports after the fact. This lack of realtime monitoring makes it all too easy to surpass emissions limits. So much so, that some producers even build noncompliance charges into their business models. Even producers with some emissions monitoring in place lack complete knowledge about how much they are emitting and when levels peak. Not only does this make compliance difficult, but it’s nearly impossible to take emissions output into account in operations and business decisions. Once oil and gas producers and the EPA have a clear grasp of actual emissions levels, the next step in the evolution of emissions regulation will be control. Everyone in the industry is watching the proposed Transport Rule, which will replace the Clean Air Interstate Rule (CAIR). The new law will set limits on emissions from utilities and lay out cap-and-trade requirements. This could foreshadow what is in store for the oil and gas industry. By anticipating and preparing for upcoming regulations now, producers can eliminate waste and make improvements to their bottom line. The ability to closely track their emissions, and tie this output information back to data based on process inputs, opens up a vast assortment of opportunities to use this information to reduce energy consumption and optimize production for more efficient operations. HYDROCARBON PROCESSING NOVEMBER 2010

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SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

A Texas-based petrochemical producer, driven by state emissions reduction requirements, implemented a PEMS to monitor and verify the continued reduction of NOx emissions by 50%. As a result, the company began to operate well within permitted NOx levels and is now able to sell (vs. buy) NOx credits through the Texas Commission on Environmental Quality Emissions Banks and Trading program. The PEMS provided detailed information on boiler output vs. input allowing the company to run its energy center’s four 300-MW boilers at their lowest oxygen levels by eliminating their artificial oxygen floor. The plant improved energy efficiency by 1%, according to a Department of Energy calculation based on temperature in the stack. Through its energy-efficiency savings alone, the plant saw full payback on the project in 13 months.

CO NOx Efficiency

Maximum efficiency

Fig. 1 illustrates the formation of NOx and CO as a function of excess O2 within a typical combustion application. Efficiency is also shown as a function of excess O2. As excess O2 decreases, NOx decreases while CO and efficiency increase. To minimize NOx production and maximize efficiency, excess O2 needs to be minimized; however, equipment safety limitations and environmental regulations limit the amount of CO that can be produced. As a result, excess O2 can only be reduced until the CO constraint is reached. At this point, combustion has been optimized, with the unit operating at the maximum allowable CO production, which corresponds to the minimum NOx production and maximum efficiency. Continuous emissions monitoring (CEM) systems.

New regulations will be the impetus for many oil and gas processors to invest in a continuous emissions monitoring system. However, when choosing which system to implement, processors should carefully consider a system that not only eases compliance, but that can also help improve productivity and reduce waste at the same time. The primary distinction between systems on the market is hardware- vs. software-based technologies. Hardware-based CEM systems. Hardware-based CEM

74

I NOVEMBER 2010 HYDROCARBON PROCESSING

NOx, efficiency

CO

systems require a significant investment in new equipment. The systems require probes to be fitted to the wall of a plant’s smokestacks to take samples of process gas at a continuous basis. These probes require weatherproof enclosures to protect them from the extreme conditions of the stack. All equipment with exposure to process gas must be designed and built for operation in hazardous locations. Additional cabinetry and shelters are required for the sampling equipment, gas analyzers and the control system that will feed emissions information back into a computer for the CO limit Minimum NOx EPA-required data collection and reporting. The two primary types of hardware-based CEM systems are direct extractive and dilution extractive. In a direct extractive sys0.00 1.00 2.00 3.00 4.00 5.00 tem, a sample pump located in a shelter fitted to a smokes stack O2, % pulls a sample of emissions through a heated probe. The probe provides filtering via a 2-μ ceramic element. The hot/wet sample NOx and CO efficiency. FIG. 1 continues through a heated sample line and onto a sample chiller located within the CEM cabinet. The sample is then flash-cooled to 4°C and water is rapidly removed. The now dry and very cold Direct extractive CEM system Analyzer sample is delivered to the analyzers on a dry Stack wall Undiluted Undiluted basis. The analyzers, in turn, provide an outhot wet dry sample put to either a datalogger or programmable Extraction sample Cooler Filter and PLC/DL probe dryer flow control logic controller (PLC). A direct-attached storage (DAS) polling computer collects this Heated probe Undiluted and sample line dry sample data and produces reports for the EPA and/ DAS Analyzer or state compliance. Hot wet (option) Analyzer In a dilution extractive CEM system, the sample is immediately mixed with or Dilution extractive CEM system Analyzer “diluted” by clean, dry and extremely cold Stack wall Diluted air (at –40°C) within the probe. ApproxiDiluted wet sample mately 5L of this air is sent up an umbilical Dilution wet sample Pneumatic PLC/DL into the probe. This air is used to develop a probe controller vacuum on a critical orifice located within Diluted the probe. This orifice allows a precise wet sample DAS Analyzer amount of process gas to blend with the Analyzer clean dry air delivered to the probe. This diluted process gas flows down the sample FIG. 2 Hardware-based CEM systems. tube in the umbilical under positive pressure. In most applications, it is not necessary


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INTERNATIONAL REFINING AND PETROCHEMICALS CONFERENCE June 2011 • Beijing, China The International Refining & Petrochemicals Conference is a market-leading program for technical and operating management in the Hydrocarbon Processing Industry. This conference will offer you an effective means to market to engineering and operations management in the HPI. Like Hydrocarbon Processing, the International Refining Conference focuses on providing the industry the very best technical content. www.GulfPub.com/IRPC

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WOMEN’S GLOBAL LEADERSHIP CONFERENCE IN ENERGY & TECHNOLOGY October 2011 • Houston, Texas Hosted by both World Oil and Hydrocarbon Processing, the Women’s Global Leadership Conference in Energy & Technology is the largest women’s event in the industry, and the only one that focuses on discussing the industry’s key environmental, economic, professional development and human capital issues in one setting. Attendees leave the conference with an increased understanding of the full range of pertinent issues and an increased ability to be change agents in our industry. This conference continues to encourage the growth and leadership of women in the industry and the respect and knowledge of energy and technology. www.WGLNetwork.com For more information about Gulf Publishing Company events or to work with us to create a new event, visit www.GulfPub.com/Events, e-mail Events@GulfPub.com, or call +1 (713) 520-4475.

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PLANT SAFETY AND ENVIRONMENT

Views

Metadata warehouse (optional)

Tx

Ti No Dif. Greater than tolerance

Model analytic engine Txv, Fyv, ... Pzv

Yes Sensor models

Cache

Predictions Tx´, Fy´, ... Pz´

Tx´

Validated sensor data

Alarms Ti´

Tx, Fy, ... Pz Raw values

Data reconciliation

Outputs NOx O2 CO Emission alarms Sensor alarms

PLC

UMS

Historians

Analyzer

Manual entry

Predictive emissions monitoring system.

to heat the sample line. In other applications, the sample dew point hasn’t been sufficiently lowered, and heating is required. The process gas is now delivered to the analyzers on a wet basis. No water was removed from the sample. It was only diluted. The analyzers then provide an output by the same method as direct extraction systems. Fig. 2 illustrates that the hardware-based approaches rely on precise instrumentation and sensors. They require high installation costs, real-estate requirements and sustained skilled maintenance. In continuous emissions monitoring and reporting, even small human errors can lead to inaccurate reporting, missed deadlines, permit violations, financial penalties and ultimately lost profit. Hardware-based approaches lack the accuracy, timeliness, consistency and security demanded by today’s regulations. Software-based CEM systems. A software-based CEM, known as a PEMS, can help oil and gas producers cost-effectively monitor emissions data without installing and maintaining expensive hardware and systems. The EPA Emissions Measurement Center has recognized predictive monitoring as a viable alternative to hardware CEM systems since 2005 and most states have followed the EPA’s guidance.* Predictive monitoring continuously monitors emissions by developing an online model using historical and real-time data from existing plant sensors. This hybrid modeling technology incorporates nonlinear empirical models such as neural networks, *The

EPA’s Emission Measurement Center developed and published a PEMS protocol and performance specifications, available on the EPA’s TTN website—http://www.epa.gov/ttn/emc/perfspec/ps-16.pdf

76

Inputs Fuel flows Fuel quality Air flow Process O2 Temperatures Ambient Humidity

Data qualification and sensor validation

DCS

FIG. 3

Browser-based client

Control console

Cache

Sensor validation

Real-time environmental management reports (optional)

Application server

SPECIALREPORT

I NOVEMBER 2010 HYDROCARBON PROCESSING

as well as first-principles models to provide the most accurate prediction models available in the industry. These models, crucial to accurate monitoring and compliance, can also be used to better understand and, therefore, control energy use. Models are executed online, using a proprietary analytic engine to provide real-time predictions of emissions from a wide range of sources and fuels. Model validation is a routine that applies known values to the sensor inputs and verifies the values against known outputs. Predetermined input values are applied to the PEMS and output values are then calculated. These values are compared to the known output values from known input values developed during modeling and relative accuracy test audits (RATAs). Values are compared and the software determines that they are in accuracy compliance. This process is an equivalency to an EPA-mandated quarterly audit. Achieving the EPA-mandated minimum operational standard for demonstrating continuous compliance with PEMS requires the ability to continue to provide reliable data in the event of sensor failure. This requires a methodology for detecting and compensating for those failures. The predictive monitoring software CEM uses a sensor validation model as a filter to detect sensor failures and to set alarms when it identifies a faulty sensor. If a failed sensor needs repair, the sensor model is also able to act as a substitute by reconstructing sensor values from the other sensor values in the plant. This allows the PEMS model to continue to accurately predict the emissions regardless of the interruption of replacing a faulty sensor and the associated downtime with the failed hardware. Fig. 4 shows predictive monitoring software compares modeled sensor output with actual sensor data for sensor validation.


PLANT SAFETY AND ENVIRONMENT

stream natural gas company was able to achieve a 90% reduction in NOx emissions, using predictive monitoring software along with ultra-low NOx burner and selective catalytic reactor (SCR) technologies at one of its natural gas plants. The plant processes up to 700 MMscf of raw natural gas per day. When its “grandfathered” operating limits for NOx emissions expired, the Texas Commission on Environmental Quality mandated that it reduced emissions by 90%. No small task. The plant installed a new boiler with ultra-low-NO x burners and a SCR. The SCR is an integral component to the boiler exhaust and it uses a specialized catalyst for the reduction of NOx emissions. The company knew that an accurate and reliable measure of NOx was critical to the feedback control loop for SCR/ ammonia injection optimization. Once this equipment was in place, the company had to choose a CEM system under considerable challenges, including: nonlinear emissions of the combined low NOx boiler with the SCR ammonia injections, inherent low-measurement signal-tonoise levels, and extreme operating conditions within the emission stack. The impact of these challenges meant the potential for penalties or fines associated with downtime as well as suboptimal economic control of ammonia usage due to inconsistent and inaccurate feedback measurements.

700 600 500 400 300 200 100 0 -100 -200 -300

Alarms

Final validated signal

Sensor model output Raw sensor measurement flagged as error

November

July

FIG. 4

CO high load

Sensor validation model output and validated signal for failed temperature transducer.

CO low load

Total stack losses at low load Heat loss, %

Clean up, comply, cut costs. A large Houston-based mid-

The producer implemented the predictive monitoring software system and achieved 6% relative accuracy in its first attempt at RATA certification, well above the EPA regulatory requirement of +/–20%. Software CEM offers hybrid modeling, through empirical models and first-principles knowledge, to provide an extremely realistic representation of process behavior. This approach, compared to other approaches that use a look-up table, helps provide improved emissions predictions even in the extreme operating ranges of the unit operations—especially for ultra-low-NOx burner boilers and turbines. Unlike historical reporting systems that provide latent information from hardware-based CEM systems, software CEM operates in real time, allowing the plant to monitor operating conditions that could affect final emissions output. This predictive methodology gives the company the ability to simultaneously incorporate process behavior and feedback into the control strategy of the boiler and SCR, making it the first of its kind to do so. In addition to helping prevent fines for noncompliance, the plant saved more than $100,000 per year in costs relative to maintaining a hardware-based CEM system. The automated, real-time and on-demand reporting capabilities of the system saved more in

Carbon monoxide CO, ppm

Armed with accurate, real-time emissions information, producers can identify vulnerabilities before they cause emission limit violations and continually factor emissions into plant automation strategies to both reduce emissions and improve efficiency. Utility, industrial/commercial/institutional (I/C/I) combustors and process heaters (P/H) are fired via a carbon-based fuel such as natural gas, oil, coal or some form of biomass that is continuously fed into the combustor chamber. During this process, complete fuel combustion occurs. However, using these fuels can be minimized by optimizing burner efficiency, thus reducing fuel consumption and emissions output. Efficiency is a direct correlation of data derived from the measurement of flue gas temperature and oxygen (O2). The combustion chamber of an industrial/commercial boiler introduces the primary air and fuel. The fuel is introduced through a burner nozzle and is designed to produce a flame front over the full range of operating conditions. Complete combustion is a function of oxygen and temperature: the greater the amount of excess oxygen, the less fuel efficient the boiler will be. Reducing the excess oxygen improves efficiency. Conversely, as oxygen is reduced, CO can start to form. Excess levels of CO indicate incomplete combustion and increased emissions. The use of CO monitoring with a PEMS can ensure that emission levels remain within limited parameters and maximize the efficiency of the burner operation. Monitoring O2 and CO as a process application for burner/boiler control has been widely accepted in the utility and industrial source market and has played a strong role in optimizing burner/boiler efficiency (Fig. 5). A PEMS is also cheaper. By eliminating the instrumentation required for a hardware-based approach, this software can predict emissions based on process variables. Across industries, a PEMS costs, on average, around 50% less than a hardware-based CEM system. In industries like oil and gas processing, where hardware must be able to withstand extremes and function in hazardous locations, installation and maintenance of a PEMS can cost 70% less than hardware solutions.

SPECIALREPORT

Total stack losses at high load

300 50 CO control band (unaffected by boiler load change) FIG. 5

Excess air Variation in O2 setpoint

Representation of the relationship of O2 and CO with efficiency. HYDROCARBON PROCESSING NOVEMBER 2010

I 77


SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

reporting costs. Over the project’s life cycle , savings will continue to grow due to higher reliability and lower maintenance costs of the software CEM. Implementing the software. Deploying a PEMS is much faster than the hardware-based alternatives. Implementing a software-based CEM usually begins with discussing project scope. Participants establish the monitoring system’s functional design and identify the necessary process sensors. If there is no existing process in place (i.e., hardware CEM system or other monitoring system), engineers obtain process data from the process’s data historian or data logger. If no emissions data is available from an existing hardware-based CEM systems, an environmental testing firm can concurrently collect stack emissions data. The emission unit goes through its entire range of operation over a two- to seven-day period (depending upon the complexity of the unit) while data is collected. Engineers use the process and stack data to construct a highly accurate emissions model. Beyond compliance. Regulatory monitoring and reporting,

followed by tighter regulatory control, intend to make oil and gas production “greener.” Processors will need to make investments in new monitoring and reporting systems to meet regulatory objectives. Installing a PEMS helps protect a plant from regulatory emissions and reporting violations at a fraction of the price of a hardware-based CEM system. It also provides plants with a mechanism to better control emissions going forward. Implementing environmentally friendly approaches to oil and gas processing and optimizing productivity need not be mutually

exclusive. By providing plant engineers with accurate predictions of emissions output in a large variety of circumstances, PEMS allows these engineers to make strategic emissions management decisions. This facilitates more efficient processing and gives plants a leg up in future GHG cap-and-trade markets. PEMS data and guidance can help oil and gas processors meet goals for cleaner air, healthier communities and a stronger bottom line. An emissions data collection is an asset that can make processing plants more competitive. A PEMS helps make that collection more accurate and affordable. HP

Richard Hovan has over 34 years of experience in the combustion, safety, environmental instrumentation and air pollution control equipment field. In addition, he has worked on grass-roots projects such as an 1,800-MW power plant in Egypt and and several water-treatment facilities. Mr. Hovan’s background includes such companies as Graver Water Company, Infilco Degrimont, Inc., Environmental Elements Corporation. He has been the vice president of product technology for KVB/Analect, vice president of combustion sales and product marketing for Land Instruments International, markets manager/corporate strategist at Forney Corporation and, most recently, manager of environmental solutions for Rockwell Automation. He has extensive overseas experience with all of his previous companies. Mr. Hovan’s knowledge includes continuous emissions monitoring (CEMS), boiler and ground-based turbine combustion efficiency monitoring, air pollution control equipment, coal pre-ignition detection systems and water treatment equipment. He has developed several new processes and has been instrumental in the development and introduction of new technologies into the market place. Mr. Hovan has served on the US-EPA’s technical review board for compliance assurance monitoring (CAM), open market trading rule (OMTR) and medical waste Incinerators (MWIs) and an opacity performance specification rewrite. He has developed the first conditional performance specification (CPS) for the EPA. He is an active member of many organizations.

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Consider switching to Internet protocol surveillance Here’s a checklist to make the jump M. S. WILSON, Infinova, Monmouth Junction, New Jersey

O

rganizations have experienced a dramatic demand for surveillance technology development to protect people, as well as private and public assets. The growing need for increasing security, especially surveillance, leaves many security managers in a quandary. A challenging dilemma that petrochemical industry security managers face is how and when they should take the leap from an analog system to an Internet protocol (IP)/digital video system. They want to jump to IP surveillance in a costmanaged way that extends the existing equipment life. For most sites, this migration will take place gradually and, during the process, analog and IP solutions will have to coexist, in some cases, for many years to come. Traditionally, in the leap from analog to digital video, organizations convert analog signals to digital signals by buying and installing rack encoders for their bank of analog cameras. They replace the analog control room equipment with new IP control room equipment. This can be quite expensive at the front end. Some believe that a better way is to create a coexistent system. In this scheme, the system keyboards connect to a virtual memory system (VMS), not the matrix switchers. The analog side of the coexisting system stays untouched. Nothing is added to it. However, since the VMS sits on top of the system, operators use their traditional keyboard commands to manage both the analog and digital solutions. This is true since the VMS interfaces with both the system’s analog matrix switchers as well as the IP cameras. As a result, on the combined video wall, the analog and IP solutions coexist but are still separate. Transparent to the operator, with no mouse needed, the system sends IP camera images

to the digital monitors and analog camera signals to the analog monitors. With this coexistent solution, agencies can begin using an IP solution simply by adding IP cameras, digital monitors and the coexistence VMS. In the leap from analog to digital, five major system areas need to be considered: • Cameras • Transmission and cabling, including power supplies • Storage and retrieval • Command and control • Integration. Cameras—throw out the analog or keep them. A key consideration for

security professionals is whether or not the existing analog cameras or new IP ones will provide the image quality needed to achieve the functional requirements of the system. Different applications have different requirements; some users require the ability to see and track suspects in changing lighting conditions, while others simply need to see that a corridor is clear. In many migration plans, specific locations of greater vulnerability or image detail requirements are ideal places for IP-based cameras, including megapixel and highdefinition models, and one needs to ask if higher-resolution cameras can help at each location (Fig. 1). A risk/vulnerability matrix can display overall elements or drill down to specific locations such as the perimeter, parking garage, entrance and exit doors, hallways, computer center, security command and control (Fig. 2). Typically, a hybrid approach is considered in which analog-to-digital encoders at the camera end can transform images from an analog camera to digital transmission and storage. The analog control room equipment gets scrapped but the new

IP control room equipment controls the already-installed analog cameras. Coexistence is a more cost-effective approach that holds down the budget at the beginning. The existing analog equipment, including cameras, control room, video wall and cabling remains untouched. VMS software, integrated with the present keyboard, sits on top of the system to manage the new IP equipment and the alreadyinstalled analog system. The petrochemical industry has another major camera challenge as well—the threat of explosives. Cameras must be in explosion-proof product housings that meet stringent corrosion-resistant requirements (Fig. 3). These housings should be made of 316 stainless steel and be suitable for both indoor and outdoor installations. The camera should comply with the IECEx standard and have an IP66 environmental protection designation. The standard for explosionproof camera housing is ExdIICT6.

FIG. 1

High-definition (HD) 36° continuous rotation megapixel IP PTZ dome cameras with 1.3-megapixel resolution.

HYDROCARBON PROCESSING NOVEMBER 2010

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SPECIALREPORT

PLANT SAFETY AND ENVIRONMENT

Likelihood assessment Reduced likelihood

Impact assessment

Reduced impact

Vulnerability

Threat experts

Vulnerability experts

Mitigation strategies

High Medium Low Criticality

Critical assessment Transportation professionals

FIG. 2

FIG. 3

Risk/vulnerability marix.

Explosion-proof camera housing.

increased bandwidth impact on the enterprise’s network. This is a tricky assignment and IT can help. Newer types of compression, decompression or codec, such as H.264, reduce bandwidth traffic load but at a cost of more storage and command center processing. Can the budget afford the increased transmission and storage associated with megapixel cameras? Storage and retrieval challenges.

There are also explosion-proof integrated pan tilt zoom (PTZ) cameras that are designed for use in flammable and explosive hazardous areas. The housings on these explosion-proof cameras are sealed to prevent the intake of explosive dust while featuring strong corrosionresistant capabilities. Transmission choices—budgets can dictate. Coaxial, shielded twisted

pair and unshielded twisted-pair cable, fiber optics and, to a lesser degree, a variety of wireless approaches carry most security video. The difference and business advantage of the various transmission schemes are in the installation and maintenance costs. A question to ask is whether or not the new IP cameras will eliminate long distance analog cabling. One strategy to handle both analog and digital networks is to transmit all the signals over a single fiber optic cable that is secure and immune to electrical or environmental interference. Installation is dramatically simplified by eliminating the need for multiple fibers, transmitters and receivers. Not to be forgotten are power supplies. Following a coexistence plan, power supplies that are multi-tap, addressable and programmable have advantages. Other considerations include the 80

I NOVEMBER 2010 HYDROCARBON PROCESSING

Though being analog-based, most security organizations already have digital and network video recorders for storage and retrieval. However, storage solutions have their own challenges, thanks to myriad features and benefits that can range from common specs to helpful elements such as intelligent PTZ control with preset positions and e-mail or SMS message notification upon motion detection or event alerts. Migrating from MJPEG to H.264 can reduce storage use by 50% or more. That’s why security users migrated from MJPEG to MPEG-4 and now are moving to H.264. It compresses video into a smaller size, yet maintains the same video quality when compared with an MPEG-4. With an H.264, a representative frame (R-frame) is selected from a group of frames in a video sequence. Only the selected R-frame is stored. By using R-frames, H.264 can compress a video stream, thereby more efficiently generating significantly less bandwidth. This is true for most camera situations, such as a fixed camera with a low amount of motion. If there is a lot of motion, as in an airport lobby, or if the camera is moving, such as a PTZ, the number of the R-frames generated will increase. In some situations, the compression provided by H.264 may be only marginally better than M-JPEG.

Nonetheless, because of the lower bandwidth generated by H.264, less storage is required to archive the video. Overall, in most surveillance situations, H.264 is a more efficient codec to use for both bandwidth reduction and storage. At the camera edge, security managers are deploying secure digital or secure digital storage cards, as well. This is especially important in applications where connection loss to the rest of the system could lead to lost images. Regardless, there are several questions to consider before selecting one mode or another on the pathway to IP: • If the video is being monitored from a remote location (and it typically is), will one get exception reporting? • Do files ever need to be shared with other departments, including law enforcement, in real time? • How much does one need to record and how long does one need to keep those recordings? Command and control options.

There is a lot to consider with command and control. Traditional matrix switching and joysticks are workhorses but in a fastapproaching software world, a solid next step is to consider networked video matrix switchers. Integration. True security systems integration is a goal of most security operations. Beyond relays and interfaces, seamless integration of security video with electronic access control, intrusion, perimeter, and identification systems is a beneficial endpoint of any operation and one made simpler through IP. The bottom line for security operations. No matter the speed of the

change-over, a solid plan is where both analog and IP cameras can coexist. Such coexistence increases security’s overall situational and domain awareness, improves its operational effectiveness and efficiencies, and provides a growth plan that extends the existing equipment’s life. It also makes systems affordable and easy to manage. HP

Mark S. Wilson has overall responsibility for Infinova’s product marketing and global marketing initiatives as well as extending relationships with manufacturing partners. He focuses on building a program structure for global marketing activities and matrixes with product management and manufacturing to develop efficient processes for new product launches and marketing operations.


ENGINEERING CASE HISTORIES

Case 59: Heat-up rates and thermal cracking A good analysis is usually better than speculation T. SOFRONAS, Consulting Engineer, Houston, Texas

A

n extruder is analyzed in this case history. While not a piece of machinery familiar to most engineers, the techniques used in this analysis can just as well be applied to pumps, compressors, heat exchangers, vessels or piping. The extruder in this example processes polymer. Its main component, a barrel, is essentially a thick pipe with longitudinal cooling passages in the wall. A screw within the barrel pushes product through a die at the discharge end. This pushing force generates a temperature rise due to the shearing action between the rotating screw and polymer. To keep the temperature constant at 525°F, water is pulsed through the drilled passages; the water turns to steam at a moderate temperature, which then helps to cool the product (the polymer) and keeps it from degrading.

Problem. A malfunctioning control valve allowed water that was not pulsed to go through the cooling passages at 150°F. There was concern that the sudden increased temperature difference could cause cracking between the water passages and the barrel bore. Figs. 1a and 1b represent a finite element model that has been reduced to ¼ symmetry and will produce results similar to a full model. The longitudinal stress is considered to be negligible, meaning it isn’t restrained axially. After five seconds, Fig. 1a shows the extruder at its 525°F product temperature, except for the cooling passage bore area, which is at 200°F. Fig. 1b shows that the stress in the bore is 80,000 lbs/in.2 in the x direction (horizontal) and also in the y direction (which is not shown here). If there were any cracks in the cooling passage, as illustrated in Fig. 2, this tensile stress would try to open them. If there was a high tensile stress region going from the bore to the cooling passage, this would be a concern since a crack could have propagated through this region. The tensile strength of the extruder barrel is 140,000 lb/in.2 It is made of ductile material, with no significant tensile stresses from the passages to the bore. For these reasons, the probability of a through-crack developing is judged to be relatively low. Applied results. The importance of this case history is to show the effect of suddenly applied temperatures and stresses they produce. Methods, such as presented here, can be used on all types of heat-up and cool-down situations. The model includes time, heat transfer coefficients and internal pressure. Of course, the analysis makes many assumptions. However, if someone were to ask if the process should now be shut down for a barrel inspection, you could provide adequate information to say that it probably has not cracked due to the malfunctioning valve.

This type of analysis obviously tests one of the possible causes and does so rather quickly. In any problem-solving session, there will be many potential causes from which the most probable will need to be addressed. All proposed causes deserve to be tested in some way. Analytical analysis is one way to evaluate them. HP

FIG. 1

a) Temperature profile at 5 sec. b) Stress profile at 5 sec.

FIG. 2

Typical temperature-induced cracking.

Dr. Anthony (Tony) Sofronas, P.E., was worldwide lead mechanical engineer for ExxonMobil before his retirement. Information on his books, seminars and consulting, as well as comments to this article, are available at http://mechanicalengineeringhelp.com.

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(4) Requested copies distributed by other mail classes

gitimate paid and/or requested distribution; (1) Out-

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ADVERTISERS in this issue of HYDROCARBON PROCESSING Company Website

Page

RS#

ACS Industries Inc. . . . . . . . . . . . . . . 59 (162) www.info.hotims.com/29425-162

ADT Commercial . . . . . . . . . . . . . . . . 26

(54)

Ametek Process Instruments . . . . . . . 28 (155) www.info.hotims.com/29425-155

Applied Analytical . . . . . . . . . . . . . . . 64 (165) (53)

www.info.hotims.com/29425-53

(80) (69)

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Borsig GmbH. . . . . . . . . . . . . . . . . . . 40 (158) (74)

www.info.hotims.com/29425-74

Cameron . . . . . . . . . . . . . . . . . . . . . . 34

(55) (67)

www.info.hotims.com/29425-67

Chemstations Inc. . . . . . . . . . . . . . . . 32 (157) www.info.hotims.com/29425-157

Costacurta SpA Vico . . . . . . . . . . . . . 53

(57)

www.info.hotims.com/29425-57

Delta Valve . . . . . . . . . . . . . . . . . . . . 22

www.info.hotims.com/29425-119

(54)

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Flexitallic LP . . . . . . . . . . . . . . . . . . . . 5 www.info.hotims.com/29425-93

www.info.hotims.com/29425-166

Newton's . . . . . . . . . . . . . . . . . . . . . 59 (163) www.info.hotims.com/29425-163

Construction Boxscore . . . . . . . . . . . 29 (156)

Ohmart/Vega . . . . . . . . . . . . . . . . . . 68 (168) www.info.hotims.com/29425-168

Events—Save the Date . . . . . . . . . . 75

Paharpur Cooling Towers, Ltd. . . . . . . . 6

www.info.hotims.com/29425-170

Webcast . . . . . . . . . . . . . . . . . . . . . 72 (169) www.info.hotims.com/29425-169

(68)

(77)

www.info.hotims.com/29425-77

Rentech Boiler System . . . . . . . . . . . . . 2

(58)

www.info.hotims.com/29425-58

Siirtec Nigi SpA . . . . . . . . . . . . . . . . . . 4 (151) www.info.hotims.com/29425-151

Spirax-Sarco Limited . . . . . . . . . . . . . 58 (161)

HPI Marketplace . . . . . . . . . . . . . 82–83 (71)

www.info.hotims.com/29425-161

Spraying Systems Co. . . . . . . . . . . . . 57

www.info.hotims.com/29425-71

Hunter Buildings . . . . . . . . . . . . . . . . 20 (153) inprocess Technology Consulting Group S.L. . . . . . . . . . . . . . . . . . . . 25 (154) www.info.hotims.com/29425-154

(62)

www.info.hotims.com/29425-62

T.D. Williamson . . . . . . . . . . . . . . . . . 87

www.info.hotims.com/29425-153

(66)

www.info.hotims.com/29425-66

Team Industrial Services. . . . . . . . . . . 21

(73)

www.info.hotims.com/29425-73

(82)

United Laboratories International,

(81)

www.info.hotims.com/29425-164

www.info.hotims.com/29425-82

Llc/Zyme-Flow . . . . . . . . . . . . . . . . 63 (164)

www.info.hotims.com/29425-81

MBI Leasing LLC . . . . . . . . . . . . . . . . 18 (152) Merichem Company . . . . . . . . . . . . . 43

(93)

Merichem Company . . . . . . . . . . . . . 45

(78)

Veolia Water Solutions & Technologies 66

(60)

www.info.hotims.com/29425-60

www.info.hotims.com/29425-78

www.info.hotims.com/29425-79

UOP LLC . . . . . . . . . . . . . . . . . . . . . . 10 Veolia Environment . . . . . . . . . . . . . . 39

www.info.hotims.com/29425-152

(88)

www.info.hotims.com/29425-88

RS#

MSA Instrument Division . . . . . . . . . . 67

Books . . . . . . . . . . . . . . . . . . . . . . . 64 (166)

Heurtey Petrochem . . . . . . . . . . . . . . 60

Page

www.info.hotims.com/29425-159

Gulf Publishing Company

Linde Process Plants . . . . . . . . . . . . . 53

Eaton Filtration . . . . . . . . . . . . . . . . . 12 (119)

Farris Engineering . . . . . . . . . . . . . . . 30

Merichem Company . . . . . . . . . . . . . 47 (159)

KBC Advanced Technologies Inc . . . . . 14 (83)

www.info.hotims.com/29425-83

Emirates . . . . . . . . . . . . . . . . . . . . . . . 8

(84)

Honeywell Analytics. . . . . . . . . . . . . . 51

www.info.hotims.com/29425-55

Chas. S. Lewis & Co., Inc. . . . . . . . . . . 48

Garlock Sealing Technologies . . . . . . . 16

www.info.hotims.com/29425-68

www.info.hotims.com/29425-158

Burckhardt Compression Ag . . . . . . . 51

Company Website

HPI Market Data Book . . . . . . . . . . . 78 (170)

www.info.hotims.com/29425-80

BJ Services . . . . . . . . . . . . . . . . . . . . 19

RS#

www.info.hotims.com/29425-156

www.info.hotims.com/29425-165

Baldor Electric Company . . . . . . . . . . 24

Page

www.info.hotims.com/29425-84

www.info.hotims.com/29425-54

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I 85


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

Utility water boot camp for process engineers—Part 3 In Part 3, we will discuss the typical responsibilities of a utility process engineer to manage their systems and to interface with process engineers at all of the production units within the complex. Plants often assign a newly graduated engineer to the utility water area. The new engineers lack of plant experience and technical knowledge of water treatment, coupled with their lack of authority over the utility water systems in the operating units create many challenges. Proactive efforts. Process-side issues demand most of the daily attention of process engineers, and utility operations are small and often-overlooked technical areas. Utilities process engineers should discuss problems with the unit process engineers. Likewise, all process engineers should read the weekly service reports from chemical suppliers; discuss nonconformances with shift supervisors and/or operators; participate in equipment inspections (boilers and heat exchangers); and become acquainted with the routine testing conducted by the operators in their unit. Sometimes, the chemical supplier will offer training for process engineers. More often, the service representative will coach the process engineer on technical issues related to their individual process unit. Table 3 summarizes typical requirements for the new process engineer assigned to utilities.

Summary. Engineers assigned to the water plant can reduce

the risk of failure by understanding the “vital few” issues that can progress rapidly and have catastrophic consequences, as well as the problems that develop slowly but have equally severe modes of failure. Slowly developing problems create a huge risk of normalization of deviance. Utility process engineers must guard against allowing this “culture of danger.” New process engineers should rely on their chemical suppliers for data interpretation and technical information about controlling corrosion, deposition and microbiological fouling. Engineers in each process unit should assume responsibility for the proper operation of the utility water systems in their unit, and collaborate with the chemical supplier and utility-water process engineer in diagnosing problems and implementing effective corrective actions. HP End of series: Part 1, September 2010, and Part 2,

October 2010. The author is president of MarTech Systems, Inc., an engineering consulting firm that provides technical services to optimize water-related systems (steam, cooling and wastewater) in refineries and petrochemical plants. She holds a BS degree in chemical engineering and is a licensed professional engineer in New Jersey and Maryland. She can be reached at: huchler@martechsystems.com.

TABLE 3. Typical utility process engineer responsibilities Action or responsibility

Frequency

Importance

Interface with service representative

• When problems occur

from chemical supplier

• When weekly reports or historical data show significant nonconformances

During annual reviews •

Minimum of once per month

Review routine operator testing

methods, data management and

• Annually

Identify key issues to correct

Limit severity of problem Learn about water treatment from chemical supplier representative

Initially when assigned to unit

monitoring procedures

corrective action procedures Review historical trend data

• When problems occur

Partner with chemical supplier to ensure proper

Confirm conformance to OEM specifications •

Limit severity of problem

• When weekly report shows nonconformances High boiler feedwater total hardness

86

Poor softener or demineralizer operation •

Condensate contamination •

Feedwater heater leak •

Feedwater pump seal leak

I NOVEMBER 2010 HYDROCARBON PROCESSING

Calcium-rich deposits will cause long-term overheat failures in boiler tubes


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