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PROCESS CONTROL AND INFORMATION SYSTEMS 速

HydrocarbonProcessing.com | OCTOBER 2012

Better control systems and equipment vastly improve operations and contribute to increased safety and reliability, along with higher profits

Our safety experts talk safety. Our operators talk control. But when it comes to keeping our people and plant safe, we all need to speak the same language.

YOU CAN DO THAT Eliminate uncertainty, reduce your risk with DeltaV SIS. Emerson’s smart safety instrumented system provides an integrated, intuitive set of engineering tools and software that enables your team to handle configuration, alarms and device health monitoring–while maintaining the systems separation required by IEC 61511 and 61508 standards. The DeltaV SIS system reduces your training and lifecycle costs by eliminating complex data-mapping and multiple databases while helping to ensure that you’re meeting safety compliance. Learn more about safety processes and best practices by downloading the Safety Lifecycle Workbook at: www.DeltaVSIS.com/workbook Select 63 at www.HydrocarbonProcessing.com/RS

OCTOBER 2012 | Volume 91 Number 10 HydrocarbonProcessing.com

SPECIAL REPORT: PROCESS CONTROL AND INFORMATION SYSTEMS

35 Update hydrocracking reactor controls for improved reliability A. G. Kern

41 Use a systematized approach of good practices in pygas hydrogenation via APC J.-M. Bader and G. Rolland

47 Why don’t we properly train control engineers? M. J. King

DEPARTMENTS

6 9 15 19 26 98 100

Brief Impact Innovations Construction Construction Boxscore Update Marketplace Advertiser index

51 Consider automated fault detection systems to improve facility reliability A. J. Szladow

COLUMNS

Reliability Equipment life extension involves upgrades

Integration Strategies Recent trends shape the future of DCS Water Management Consider software tools for water reuse projects

55 Optimize desulfurization of gasoline via advanced process control techniques V. Yadav, P. Dube, H. Shah and S. Debnath

REFINING DEVELOPMENTS

61 Maximize diesel production in an FCC-centered refinery, Part 2 P. K. Niccum

SULFUR—SUPPLEMENT

S-69 Optimize sulfur recovery from dilute H2S sources M. P. Heisel and A. F. Slavens

HEAT TRANSFER

83 Identify and control excess air from process heaters S. Ahamad and R. Vallavanatt

ROTATING EQUIPMENT

91 Apply new pump-drive software to test performance K. Bihler, D. Dominiak, B. Keith and J. Johnson Cover Image: Photo courtesy of Emerson Process Management

102

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EDITORIAL Editor Reliability/Equipment Editor Process Editor Technical Editor Online Editor Associate Editor Contributing Editor Contributing Editor Contributing Editor

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Hydrocarbon Processing is indexed by Applied Science & Technology Index, by Chemical Abstracts and by Engineering Index Inc. Microfilm copies available through University Microfilms, International, Ann Arbor, Mich. The full text of Hydrocarbon Processing is also available in electronic versions of the Business Periodicals Index.

ARTICLE REPRINTS If you would like to have a recent article reprinted for an upcoming conference or for use as a marketing tool, contact Foster Printing Company for a price quote. Articles are reprinted on quality stock with advertisements removed; options are available for covers and turnaround times. Our minimum order is a quantity of 100. For more information about article reprints, call Rhonda Brown with Foster Printing Company at +1 (866) 879-9144 ext 194 or e-mail rhondab@FosterPrinting.com. Hydrocarbon Processing (ISSN 0018-8190) is published monthly by Gulf Publishing Company, 2 Greenway Plaza, Suite 1020, Houston, Texas 77046. Periodicals postage paid at Houston, Texas, and at additional mailing office. POSTMASTER: Send address changes to Hydrocarbon Processing, P.O. Box 2608, Houston, Texas 77252. Copyright © 2012 by Gulf Publishing Company. All rights reserved. Permission is granted by the copyright owner to libraries and others registered with the Copyright Clearance Center (CCC) to photocopy any articles herein for the base fee of $3 per copy per page. Payment should be sent directly to the CCC, 21 Congress St., Salem, Mass. 01970. Copying for other than personal or internal reference use without express permission is prohibited. Requests for special permission or bulk orders should be addressed to the Editor. ISSN 0018-8190/01.

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President/CEO Vice President Vice President, Production Business Finance Manager

John Royall Ron Higgins Sheryl Stone Pamela Harvey

Part of Euromoney Institutional Investor PLC. Other energy group titles include: World Oil and Petroleum Economist Publication Agreement Number 40034765

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| Brief New fuel standards in the US The US government recently finalized standards that will increase fuel economy to the equivalent of 54.5 mpg for cars and light-duty trucks by model year 2025. When combined with previous standards, this move will nearly double the fuel efficiency of those vehicles compared to new vehicles currently on the road. In total, the program to improve fuel economy is expected to reduce US oil consumption by 12 billion barrels. The standards issued by the US Department of Transportation (DOT) and the US Environmental Protection Agency (EPA) build on the previously issued standards for cars and light trucks for model years 2011â&#x20AC;&#x201C;2016. Those standards raised average fuel efficiency by 2016 to the equivalent of 35.5 mpg.

BILLY THINNES, TECHNICAL EDITOR / Billy.Thinnes@HydrocarbonProcessing.com

Brief Valero has decided to further reduce operations and reorganize its 235,000-bpd Aruba refinery as a

refined products terminal, the company said. The terminal will feature both deepwater berths and smaller berths, and will have the flexibility to load the very largest crude ships. Terminal activities will, however, require a considerably smaller workforce, according to the company. The reorganization and reduction in workforce is expected to be complete before the end of 2012. Valero will continue to supply jet fuel, gasoline, diesel and fuel oil to the island, as well as engage in third-party terminal services. In the terminal operations mode, Valero will continue to invest in Aruba with facility improvements and dock and tankage upgrades, the company said. In the nearterm, the refinery will continue to be maintained in a state that would allow a restart, should Valero be successful in the pursuit of alternatives for the refinery prior to the terminal transition. US investment group Carlyle has agreed to buy the performance coatings business of DuPont for $4.9 billion

in cash. The transaction is expected to close in the first quarter of 2013, subject to customary closing conditions and regulatory approvals. DuPont Performance Coatings is a global supplier of vehicle and industrial coating systems, with 2012 expected sales of more than $4 billion and more than 11,000 employees. As part of the transaction, Carlyle will assume $250 million of DuPont’s unfunded pension liabilities. Carlyle’s industrial and automotive investments include Allison Transmission, Hertz and PQ Corp., as well as recent commitments to invest in Hamilton Sundstrand Industrial and regional rail freight operator Genesee & Wyoming. Enterprise Products recently began an open commitment period to determine additional shipper

demand for capacity on its Appalachia-to-Texas (ATEX Express) ethane pipeline. The 1,230-mile system will deliver growing ethane production from the Marcellus/Utica Shale areas of Pennsylvania, West Virginia and Ohio to Mont Belvieu, Texas. The open commitment period will be used to determine market interest in executing additional 15-year binding transportation agreements. The ATEX Express is expected to begin operations in the first quarter of 2014. Cosmo Oil will permanently close its 140,000-bpd Sakaide refinery in western Japan by July 2013 to meet

a government regulation that encourages refining capacity cuts amid falling local demand. The Japanese Ministry of Economy, Trade and Industry set rules in July 2010 requiring refiners to raise residual cracking capacity to a designated percentage of crude refining capacity, as calculated by a formula, by March 2014. By closing the refinery, Cosmo expects to save Y10 billion a year in costs.

Technip has completed the acquisition of the Stone & Webster process technologies and associated oil and

gas engineering capabilities. Technip sees this acquisition as a way to further diversify its onshore/offshore segment, adding revenues based on technology supply. It will also use the Stone & Webster brand to expand in promising growth areas such as the US, where downstream markets will benefit from the supply of unconventional gas. To make the most of these strengths, a new business unit, Technip Stone & Webster Process Technology, will be developed within the company’s onshore/offshore segment. Technip paid cash consideration of around €225 million from existing cash resources, which will be subject to customary price adjustments. Air Liquide has officially opened its first public hydrogen filling station for passenger cars in Düsseldorf, Germany.

This station will be followed by 10 new hydrogen filling stations that will be designed, built and rolled out in the next three years under the auspices of the German government’s major demonstration project. By 2015, Germany will have a supply network of at least 50 public hydrogen filling stations. Driven by the same dynamic, two other stations have been installed recently by Air Liquide in Oslo, Norway, and in Brugg, Switzerland. In Japan, the government sees hydrogen as a promising major energy source for cars and expects to install about 100 hydrogen distribution stations for fuel cell vehicles by 2015. In response to this government policy, Air Liquide Japan has recently set up a specialized team focused on the hydrogen business. So far, they have installed three hydrogen energy stations (in Tokyo, Kawasaki and Saga). One of these stations demonstrated the feasibility of a complete “blue hydrogen” chain, from wood chips to clean mobility. Western Refining and Glencore International announced that two of their subsidiary companies (York River

Fuels and Glencore Ltd.) have entered into a long-term commercial supply and trading agreement. Glencore has agreed to provide global sourcing, supply and trading and inventory and risk management services to support York River’s midAtlantic wholesale business. In return, York River has agreed to provide rack marketing and contract and credit management. Glencore has entered into a long-term commitment with Epic Terminals at its terminal in Savannah, Georgia. The Savannah terminal includes over 450,000 bbl of storage capacity for various grades of gasoline, distillates, ethanol, biofuels and fuel blends. The terminal will enable the two companies to expand their wholesale capabilities and provide fuel products to their customers from southern Georgia to northern Maryland. Western Refining operates refineries in El Paso, Texas, and Gallup, New Mexico. The company also runs products terminals in Albuquerque and Bloomfield, New Mexico. Hydrocarbon Processing | OCTOBER 2012 7

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Impact

Increased coal-to-olefins processes in China China’s significant domestic supply of coal, combined with a domestic shortage of several key chemical feedstocks, especially ethylene and propylene, are driving increased Chinese demand for more production of chemical feedstocks from coal, according to a new IHS study that assessed the key technologies and economics of coal-to-olefins (CTO) processes employed in China. The study noted that, in 2011, China had an ethylene capacity of 15.7 MMt and production of 14.4 MMt. On the demand side, China’s total ethylene equivalent consumption (including imports of first-order derivatives such as polyethylene) far exceeded its domestic ethylene supply. China imported nearly 8 million tons of polyethylene alone in 2011, accounting for 42% of total Chinese demand. In a new five-year plan covering 2011–2015, the Chinese set a target that 20% of the country’s ethylene production will come from other diversified sources, which for China—a country with abundant coal supplies that is a netimporter of oil—practically means coal. According to IHS, China’s domestic demand for oil was 9.4 million bpd in 2011, of which 57% was imported. Likewise, China’s propylene production was 13.1 MMt in 2011. On the demand side, China’s total propylene equivalent consumption, including imports of first-order derivatives such as polypropylene, also far exceeded its domestic propylene supply. China imported nearly 5 MMt of polypropylene alone in 2011, accounting for 30% of total demand. The propylene shortage in China is projected to stay at about 5 MMtpy until 2020. The processes studied included the gasification of bituminous coal by GE Texaco or Shell gasifiers to produce synthetic gases (syngas), followed by methanol synthesis and methanol-to-

olefins (MTO) production. The MTO technologies studied included UOP/ Hydro MTO and Lurgi methanol-topropylene (MTP) technologies. Economic evaluations were based on a US Gulf Coast location. However, since most coal-based olefin projects are occurring in China, the economics in the review were adjusted to reflect production and capital costs for a Chinese location. The adjustment was achieved by examining the variations in technologies deployed in China and accounts for capital investment, raw materials, utility and labor costs relative to the design basis used in the report. To address the country’s chemical feedstock shortage, China has built or is planning many high-capacity, integrated CTO and coal-to-propylene (CTP) plants. Thirteen plants are in the works, with four of those currently operational. According to the IHS review, all coalbased processes analyzed in the review showed lower direct costs, but higher indirect costs (due to high capital investments) as compared to competing (petroleum-based) processes for CTO and CTP, respectively. To enable baseline comparisons of chemical engineering processes for this review, return on investment (ROI) was the primary factor considered, and the costs were not weighted for environmental impact. For olefins production, based on the market price of olefins at the time of analysis, the MTO process based on outsourced methanol offers the highest ROI, followed by the integrated GE/ MTO process, and finally, steam-cracking of naphtha, which is a petroleumbased process. In terms of propylene production, based on the market price of propylene at the time of analysis, the MTP process based on outsourced methanol offers the highest ROI, followed by the integrated Shell/MTP process using bituminous coal, the integrated Shell/MTP using lignite, and finally, the integrated Siemens/MTP.

New nanoscale reference material to be known as P25 The National Institute of Standards and Technology (NIST) has issued a new nanoscale reference material for use in a wide range of environmental, health and safety studies of industrial nanomaterials. The new NIST reference material is a sample of commercial titanium dioxide powder commonly known as “P25.” NIST standard reference materials (SRMs) are typically samples of industrially or clinically important materials that have been carefully analyzed by NIST. They are provided with certified values for certain key properties so that they can be used in experiments as a known reference point. Nanoscale titanium dioxide powder may well be the most widely manufactured and used nanomaterial in the world, and, not coincidentally, it is also one of the most widely studied (FIG. 1). In the form of larger particles, titanium dioxide is a common white pigment. As nanoscale particles, the material is widely used as a photocatalyst, a sterilizing agent and an ultraviolet blocker (in sunscreen lotions, for example).

FIG. 1. The nanoscale crystalline structure of titanium dioxide in NIST SRM 1898 (color added for clarity). Hydrocarbon Processing | OCTOBER 2012 9

Impact “Titanium dioxide is not considered highly toxic and, in fact, we don’t certify its toxicity,” said NIST chemist Vincent Hackley. “But it’s a representative industrial nanopowder that you could include in an environmental or toxicity study. It’s important in such research to include measurements that characterize the nanomaterial you’re studying—properties like morphology, surface area and elemental composition. We’re providing a known benchmark.”

The new titanium dioxide reference material is a mixed phase, nanocrystalline form of the chemical in a dry powder. To assist in its proper use, NIST has developed protocols for properly preparing samples for environmental or toxicological studies. The new SRM also is particularly well suited for use in calibrating and testing analytical instruments that measure specific surface area of nanomaterials by the

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widely used Brunauer-Emmett-Teller gas sorption method.

Marginal increase forecast in North American lubricant market Although it is estimated that there will be a 3% increase in tonnage carried by private fleet operators in the US through 2016, this is expected to translate to a marginal increase in commercial lubricant consumption according to a new lubricants study from Kline and Company. On-highway activity saw a surge in the latter half of 2010 that continued well into 2011. Similarly, the lackluster performance of the construction industry between 2008 and 2010 has begun to show signs of a rebound. However, increased service implementation of longer drain interval oils due to a higher penetration of synthetics, growth in oil analysis practices, and an overall increase in commercial vehicles’ mechanical efficiency, mean that commercial lubricant consumption is expected to fractionally increase by a compound annual growth rate of just 0.4% to 1.0% to 2016. Shell remains the leading supplier of lubricants in North America and accounts for an estimated 12% of the market share in 2011, followed by ExxonMobil, Chevron and BP. With the growing realization of the benefits of synthetics and their consequent steady uptake, value is rising, while overall demand is being suppressed through inherently longer service intervals. Similarly, oil analysis—the laboratory analysis of a lubricant’s properties, suspended contaminants and wear debris—is being increasingly performed during routine preventive maintenance to provide meaningful and accurate information on lubricant and machine condition. By tracking oil analysis sample results over the life of a vehicle, lubricant consumption is optimized. Re-refined engine oils are slowly making their way into the commercial automotive segment; however, a majority of respondents participating in a survey for the research cited concerns about OEM approvals of such grades and the possible non-availability on the highways, as major deterrents. In particular, the US commercial trucking industry generally appears not yet prepared to accept re-refined oils;

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Impact with a majority of equipment/maintenance managers interviewed conceding that reliability and logistics issues are prime considerations and impediments. A number of farmers and farm cooperatives interviewed for this study showed minimal interest in using re-refined oils, believing that lubricants made out of re-refined basestocks are of an inferior quality. However, Tushar Raval, director of Kline’s energy practice, notes the opportunity by saying, “An immediate connect can be made by the way of marketing re-refined oils as ‘sustainable’

FIG. 2. Natural gas vehicles, like this Honda Civic, are starting to gain traction in the US.

products and consequently more easily find favor from the farming community. “Another way of successfully propagating the acceptance of these grades is by way of approvals and recommendations from OEMs, such as John Deere,” Mr. Raval said.

Natural gas vehicles could be gaining traction

A new report from PIRA Energy Group says that the sheer volume of US recoverable gas resources relative to expected demand suggests that benchmark Henry Hub gas prices will remain deeply discounted relative to oil prices beyond this decade. Furthermore, the lengthy period of low-cost gas relative to oil has tremendously broadened support for the view that inexpensive North American gas is here to stay. According to the report, by employing off-the-shelf technologies, consumers could be able to accrue substantial savings given the latent expected price advantages of natural gas vs. diesel. Such savings can also be attained in the transportation sector, particularly

with regard to the much discussed development of natural gas vehicles (NGVs) (FIG. 2). The report concludes that future gas demand in such NGVs has enormous upside potential, led by private sector initiatives, with or without federal government assistance. Adoption of natural gas into both US commercial trucking and all varieties of fleets is approaching a critical threshold, which ultimately could lead to enormous gas demand growth at the expense of diesel fuel. In an overall high case scenario, NGV gas demand would be capable of reaching 14 Bcfd by 2030, suggesting that as much as 2.4 MMbpd of diesel fuel demand could be at risk. Liquefied natural gas (LNG) consumed in Class 8 trucks would be responsible for approximately 70% of that total, 10 Bcfd. Fleet vehicles typically consuming compressed natural gas (CNG) would account for the additional 4 Bcfd. PIRA forecasts natural gas will capture a more moderate, but also impressive, 7 Bcfd share of the US on-highway transportation fuels market by 2030.

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ADRIENNE BLUME, PROCESS EDITOR Adrienne.Blume@HydrocarbonProcessing.com

Innovations Invensys introduces new technology offerings Software and technology provider Invensys Operations Management recently released a program to help clients modernize and improve the performance of aging control systems and other plant equipment. The program guides clients in calculating modernization costs, reducing risk, deploying advanced technology, and approaching plant upgrades strategically and systematically. Under the program, Invensys will deliver full-scope consulting, project management, engineering, installation and maintenance services, and products and solutions that minimize the risk of operating obsolete technologies. Invensys starts with an assessment to understand the company’s business initiatives and issues. The input received is used to develop a strategic plan that meets the plant’s business and technology needs. As part of the assessment, Invensys also helps clients establish return-on-investment targets. The company’s hardware and software offerings address all operational areas of the plant, including instrumentation, input/output (I/O) and human/machine interface (HMI), safety and critical control systems, turbomachinery assets, process safety lifecycle components, cyber security systems and other assets. In another development, Invensys has extended its virtualization technology offerings. Initially focused on the Microsoft HyperV and VMware platforms within its software product lines, the new Invensys offering now includes thin client support and intelligent solutions for the company’s Foxboro I/A series distributed control system (FIG. 1). Intended to lower total cost of ownership and promote successful project delivery, the new offerings will help customers cut implementation costs, reduce risks, shorten project schedules, improve scheduling integrity, strengthen the ability to respond to project changes, and improve global collaboration.

The hardware offerings are formulated to maximize the advantages of virtualization technology. Along with intelligent marshalling and engineering services, the offerings include a new range of servers specifically selected and qualified as an optimized virtual machine-hosting appliance, a new range of solid-state operator client terminals, thin client management software, a USB modular alarm annunciator keyboard, virtual machine-hosting software, recommendations on cybersecurity best practices, guest operating system licenses, and specialized support for Invensys control and safety offerings.

and timed events on chromatograms following instrumental variation over time. Select 2 at www.HydrocarbonProcessing.com/RS

Largest German cellulosic ethanol plant starts up Swiss specialty chemicals company Clariant recently inaugurated Germany’s biggest pilot plant (FIG. 3) in Straubing, Bavaria. The €28 million (MM) plant, which is based on Clariant’s sunliquid

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Integrated chromatography aids fuel producers Bruker’s Chemical and Applied Markets (CAM) division’s CompassCDS data handling system (FIG. 2) networks gas chromatograph (GC) instruments into closed-loop information systems in industrial and applied environments. The system is capable of interfacing with multiple middleware systems, such as supervisory control and data acquisition (SCADA) systems and laboratory information management systems (LIMS). This ensures an unbroken flow of information and rapid feedback of analytical results to support optimum processing and product validation. Additionally, by removing the need for human intervention, the risk for errors is reduced. The CompassCDS product is built around a central administrative core, known as the “configuration manager.” There are several customized modules that are specific to the petrochemicals industry, including simulated distillation, hydrocarbon analysis and PIONA+ (paraffins, isoparaffins, olefins, naphthenes and aromatics, plus oxygenates). A simple graphical user interface allows all operations to be carried out from only two screens, which enables ease of use and expedites training. Additionally, the system’s added IntelliUpdate feature automatically corrects both retention time

FIG. 1. Invensys’ virtualization system allows the user to consolidate many PCs and servers into a high-availability virtual host server.

FIG. 2. Bruker CAM’s CompassCDS system networks gas chromatograph instruments with infosystems. Hydrocarbon Processing | OCTOBER 2012 15

Innovations technology, will produce up to 1,000 metric tons of cellulosic ethanol from around 4,500 metric tons of wheat straw. The pilot plant will confirm the technological feasibility of the sunliquid technique, and the process will later be used in an industrial-scale plant. According to studies, Germany potentially has around 22 MM metric tons of straw that could be used for energy production, which would be sufficient to cover around 25% of the country’s current gasoline requirements. German Federal Minister Annette Schavan commented, “This plant clearly demonstrates that products traditionally based on petroleum can be manufactured to the same standard using biomass. This new plant serves as an important contribution to a sustainable bioeconomy.” Select 3 at www.HydrocarbonProcessing.com/RS

Software protects against cyber attacks Honeywell’s Management of Change (MOC) is an integrated, add-in module that runs on top of DOC4000 assessment software and leverages Web 2.0 technologies to facilitate information flow and collaboration. MOC enables increased safety and compliance, and it helps protect against threats of cyber attacks and safety

FIG. 3. Clariant’s cellulosic ethanol pilot plant in Straubing is the largest in Germany.

hazards at plants, by effectively managing changes and approvals. FIG. 4 illustrates the workflow process for MOC. MOC also enables improved handling of critical issues, including undocumented changes, enhanced regulatory compliance, reductions in error-prone manual MOC tasks, and unauthorized changes that increase risk. The module is specifically designed to automatically detect all automation changes, to reconcile MOC cases to changes in the automation system, and to automatically generate reports of unreconciled changes, all at a 45% lower cost than manual methods. This enables rapid root-cause analysis to ensure business continuity, and it can reduce potentially significant financial impact on production. Select 4 at www.HydrocarbonProcessing.com/RS

Technology enables ethanol production breakthrough As countries and companies evaluate their supply options to meet growing transportation fuels demand, they will need to balance four priorities: safe and clean fuel blendstocks, cost, energy security, and global environmental impact. Celanese believes that ethanol produced using Celanese TCX Technology (FIG. 5) is the best fuel choice to meet these considerations. Ethanol has already gained acceptance in most global markets as a high-octane, nontoxic, biodegradable fuel. However, traditional production processes are not economically viable, as they compete for arable land and typically require government mandates or subsidies. Celanese TCX Technology produces ethanol in a commercially viable, lowcost manner, from locally available hydrocarbon resources such as natural gas and coal, rather than from corn or sugar cane. For these reasons, no arable land use or government support is required. Select 5 at www.HydrocarbonProcessing.com/RS

FIG. 4. Honeywell’s MOC software helps protect plants against safety threats.

FIG. 5. Celanese’s TCX Technology produces ethanol from hydrocarbons.

16 OCTOBER 2012 | HydrocarbonProcessing.com

FIG. 6. Yokogawa’s STARDOM FCN controller has been certified for use as a flow computer.

Yokogawa’s controller certified as flow computer Yokogawa Electric Corp.’s STARDOM FCN autonomous controller (FIG. 6) was recently certified by Measurement Canada for use as a flow computer. The certification is based on the determination that the controller has the same accuracy as a conventional flow computer. Devices subject to its approval are used to measure gas, electricity, mass and volume, and they are tested on a broad range of criteria including design, construction, marking, accuracy and sealing method. While conventional flow computers meet all metering requirements, there is now a trend toward embedding this functionality in programmable logic controllers (PLCs) and remote telemetry units (RTUs), which are valued for their durable construction and versatile control capabilities. In addition to conventional analog transmitter signals, which can be affected by noise and variations in the ambient temperature, the STARDOM FCN controller supports field digital communication protocols such as HART, Modbus, and Foundation fieldbus for use with a wide range of transmitters. Yokogawa’s plant resource manager (PRM) asset management system increases maintainability while reducing both engineering time and the cost of monitoring widely distributed facilities. Select 6 at www.HydrocarbonProcessing.com/RS

SPECTRO wins ACHEMA Innovation Award At the ACHEMA chemical engineering and biotechnology trade show in Frankfurt, Germany, in June, SPECTRO Analytical Instruments received the Innovation Award for its SPECTROBLUE ICP-OES spectrometer (FIG. 7). Introduced in 2011, the SPECTROBLUE ICP-OES spectrometer is targeted for environmental laboratories in need of quick and accurate analyses of water, wastewater, sewage sludge and soil samples for toxic heavy metals. SPECTROBLUE’s air-cooled, optical plasma interface diverts heat away with an air stream, marking an advance in spectrometer technology. Another innovative feature of SPECTROBLUE is its improved sample introduction. SPECTRO has significantly

Innovations shortened the path of the sample into the plasma, which decreases the duration of the analysis and reduces carryover effects. SPECTROBLUE’s operating software also includes new, user-friendly functions, such as a comprehensive Smart Analyzer Vision software package and a Smart User Interface that simplifies routine operation. Select 7 at www.HydrocarbonProcessing.com/RS

Mobile tool enables portable pH reading Sensorex has developed a mobile accessory for pH measurements that is compatible with Apple iPod, iPhone and iPad devices. The patent-pending PH-1 pH meter accessory (FIG. 8) measures and records pH values in the lab or field for use in environmental, educational and industrial applications. The PH-1 accessory plugs into the standard Apple dock connector and is powered from the Apple device, requiring no supplemental energy source. It uses a Sensorex pH electrode to measure pH in a range of 0–14, with accuracy to 0.01 pH. It operates in ambient temperatures of 0°C–40°C and in solutions of 0°C–100°C. The free Sensorex app displays pH, millivolts, ambient temperature and solution temperature in real time. The CEmarked device supports one, two, three or more calibration points, and it sends readings by email for later analysis. Also, when used with a GPS-enabled device, the pH meter application will record measurements with both timestamp and

geographic coordinates, eliminating transcription errors and improving efficiency. Select 8 at www.HydrocarbonProcessing.com/RS

French consortium eyes BioOil upgrading Axens, IFP Energies nouvelles (IFPEn) and Dynamotive recently announced completion of agreements for the development, industrialization and commercialization of a proprietary process to produce transportation fuels from Dynamotive’s BioOil pyrolysis oil. The process is said to have competitive advantages compared to existing processes and competing technologies. Dynamotive will provide pyrolysis oil to IFPEn for the development program, while Axens will lead the development, industrialization and commercialization of the upgrading technology. Laboratoryscale units have been developed and operated in Canada and at IFPEn facilities in Lyon, France, where Dynamotive’s BioOil was upgraded to synthetic hydrocarbons. Dynamotive’s BioOil technology is based on the application of fast pyroly-

sis (burning without oxygen) to biomass waste (agricultural and forestry) to produce a high-quality, versatile and economic biofuel. BioOil can be further converted into vehicle fuels and chemicals. Select 9 at www.HydrocarbonProcessing.com/RS

FIG. 8. Sensorex’s mobile accessory for pH measurements is compatible with a range of Apple devices.

Recruiting “A-level” candidates for your C-level positions (Management & Sales positions also) Providing executive recruiting services to the energy markets.

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For a confidential C-level executive search or placement of management or sales positions, please call Thomas Brinsko or Raul Hernandez in Houston at 281-538-9996 or visit www.bicrecruiting.com.

FIG. 7. SPECTRO Analytical Instruments was awarded at ACHEMA for its SPECTROBLUE spectrometer.

For more information on strategic marketing through BIC Alliance, investment banking services through IVS Investment Banking or custom books, event planning or speaker services through BIC Media Solutions, contact Earl Heard or Thomas Brinsko at (800) 460-4242, or visit www.bicalliance.com.

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Breaking through Heavy Oil Barriers Facing a labyrinth of ever-changing feedstock and production demands? Invensys Operations Management helps you break through the barriers with our unique Heavy Oils models, just one part of the SimSci-Esscor suite of reďŹ nery wide optimization software solutions. For more information visit us at: iom.invensys.com/heavyoils Select 69 at www.HydrocarbonProcessing.com/RS

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ÂŠ Copyright 2012. All rights reserved. Invensys, the Invensys logo, Avantis, Eurotherm, Foxboro, IMServ, InFusion, Skelta, SimSci-Esscor, Triconex and Wonderware are trademarks of Invensys plc, its subsidiaries or affiliates. All other brands and product names may be trademarks of their respective owners.

HELEN MECHE, ASSOCIATE EDITOR Helen.Meche@HydrocarbonProcessing.com

Construction

North America Jacobs Engineering Group Inc. has a contract from Methanex Corp. to provide engineering, procurement and construction services for a methanol production facility in Louisiana. Officials estimate the construction value to be $550 million. Jacobs is already executing site-specific engineering and construction management for the 225-acre location in Geismar, Louisiana, from its offices in Baton Rouge, with support for the disassembly from its Santiago, Chile, office. The plant is expected to be operational in the second half of 2014. KBR has a general works contract for phase-two construction at a raw gasprocessing and compression facility near Dawson Creek, British Columbia, Canada. KBR’s Canadian subsidiary, KBR Wabi, will execute construction and related site support for the facility’s expansion, increasing the existing capacity to 100 million scfd. The award follows KBR’s recent work—delivering pipe-rack fabrication and module assembly for phase one of the Dawson Creek plant. A joint-development agreement, focusing on bio-based butadiene, has been signed by INVISTA and LanzaTech to develop one-step and two-step technologies for converting industrial waste-gas carbonmonoxide (CO) into butadiene. Initial commercialization is expected in 2016. Initially, the focus will be on producing butadiene in a two-step process from LanzaTech CO-derived 2,3-butanediol (2,3 BDO). A direct single-step process will also be developed to produce butadiene directly through a gas-fermentation process. INVISTA and LanzaTech will also jointly collaborate on developing tools that will extend this technology—once developed—to directly produce other industrial chemicals. These include nylon intermediates, from CO containing waste gases, using LanzaTech’s gas-fer-

mentation technology and proprietary biochemical platform. INVISTA is building internal biotechnical capability to develop biological routes to its products and feedstocks. Praxair, Inc., has broken ground on its new air-separation unit in Memphis, Tennessee. With a capacity of 600 tpd, the new plant is scheduled to start up in the second quarter of 2013. INVISTA has selected its production facility in Orange, Texas, as the initial location to install its next-generation adiponitrile (ADN) technology. ADN is a critical intermediate chemical used in the manufacture of nylon 6,6. The project to convert the Orange site to the new technology is well underway, and INVISTA is expected to invest more than $100 million at the facility in the next 18 months. The technology, a new butadienebased chemistry, is said to improve product yields and ease of operations, while requiring a lower annual-maintenance investment compared to existing technology. Evidenced through operation of a pilot-scale facility, also located in Orange, the technology also delivers significant air emission and waste reductions. The company hopes to be in full production by mid-2014. Cheniere Energy Partners, L.P. has completed all milestones and has issued Bechtel Oil, Gas and Chemicals, Inc., with a full notice to proceed on construction of the Sabine Pass Liquefaction Project’s first two liquefaction trains. The first liquefaction train is expected to start operations as early as 2015. The second liquefaction train is expected to commence operations six to nine months after the first train’s startup. Flint Hills Resources is considering spending more than $250 million to enable its West refinery in Corpus Christi, Texas, to process more Eagle Ford crude

oil, while extending its ability to reduce criteria air emissions. The company operates two Corpus Christi refineries: the West refinery, with a capacity of about 230,000 bpd, and the East refinery, with a capacity of about 70,000 bpd. Flint Hills Resources expects to submit the permit applications to the Texas Commission on Environmental Quality and the US Environmental Protection Agency in the coming weeks.

South America A subsidiary of Foster Wheeler AG’s Global Engineering and Construction Group has a contract from Petrobras for a world-scale grassroots gas-to-chemicals complex—Complexo Gás-Químico UFN-IV—in Linhares, Espirito Santo State, Brazil. Foster Wheeler will provide basic engineering design (BED), front-end engineering and design (FEED), as well as technical assistance and training during the engineering, procurement and construction (EPC) phase through to successful completion of plant performance tests. The BED and FEED will be included in the company’s third-quarter 2012 bookings. The provision of technical assistance and training will be booked at a later date, after the FEED is complete, when Petrobras advises that it is proceeding with the project’s EPC phase. TREND ANALYSIS FORECASTING Hydrocarbon Processing maintains an extensive database of historical HPI project information. The Construction Boxscore Database is a 45-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 commadelimited or Excel® and can be custom sorted to suit your needs. The cost depends on the size and complexity of the sort requested. You can focus on a narrow request, such as the history of a particular type of project, or you can obtain the entire 45-year Boxscore database or portions thereof. Simply send a clear description of the data needed and receive a prompt cost quotation. Contact Lee Nichols at 713-525-4626 or Lee.Nichols@GulfPub.com

Hydrocarbon Processing | OCTOBER 2012 19

Construction Foster Wheeler will act as integrator for the entire complex, managing the overall BED and FEED, including managing the process licensors and Brazilian subcontractors. The BED/FEED phase is scheduled for completion at the end of 2013. The complex is expected to produce in excess of 1 million tpy of ammonia and urea fertilizers, methanol, acetic acid, plus formic acid and melamine, helping to reduce Brazil’s imports of these products.

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Clariant has inaugurated what is said to be Germany’s biggest pilot plant for producing climate-friendly cellulose ethanol from agricultural waste. Located in Straubing, Bavaria, and supported by the Bavarian government and the Federal Ministry for Education and Research, the futuristic project will produce up to 1,000 tons of cellulose ethanol from around 4,500 tons of wheat straw based on Clariant’s sunliquid technology. It represents an investment of around €28 million. The sunliquid process is an innovative biotechnological method that turns plant waste products, such as grain straw and corn straw, into second-generation cellulose ethanol. Marquard & Bahls, through its subsidiary Bomin, and The Linde Group, will establish a joint-venture ( JV) company to build infrastructure for liquefied natural gas (LNG) in Europe’s maritime sector. The transaction is subject to the approval of the relevant antitrust authorities. The 50/50 JV is due to start its operations in the latter part of 2012, with its headquarters based in Hamburg, Germany. The JV will set out to establish an LNG supply chain and to provide reliable, safe and environmentally friendly fuel to ship owners and operators. Linde will contribute its vast experience in cryogenics and its best-in-class engineering know-how, while Bomin will support the JV with its excellent track record in maritime bunker-fuel trading and operations. The new company will establish operations in a number of key ports throughout the so-called “emission control areas” in Northwest Europe. CB&I has an award from BASF for the engineering, procurement and construction management of a new butadieneextraction plant in Antwerp, Belgium.

The contract, which is valued in excess of $50 million, is an essential part of the total BASF investment amount, which will be in the high double-digit million euro range. The plant is scheduled to start up during 2014. Jacobs Engineering Group Inc. has a five-year enterprise frame agreement (EFA) from Shell Global Solutions International B.V. to provide engineering and project management services to Shell’s European downstream assets. The contract has the options to be renewed for an additional five years and/or to be extended to other Shell businesses such as upstream, and beyond Europe to the Middle East and Africa. Under the EFA, Jacobs will provide services ranging from feasibility studies and small plant modifications to discrete projects for Shell’s major refining and chemical sites in Pernis, The Netherlands, and in Rhineland, Germany. UOP LLC, a Honeywell company, has been selected by Lukoil to provide technology to produce blending components used to make high-octane gasoline and petrochemicals at Lukoil’s facility in Nizhny Novgorod, Russia. Lukoil will license an integrated suite of Honeywell’s UOP technologies to produce high-quality gasoline-blending components, propylene and other petrochemicals. The new units, expected to start up in 2015, will produce more than 1 million metric tpy of gasoline-blending components and more than 170,000 metric tpy of propylene. In addition to technology licensing, Honeywell’s UOP and a number of its affiliates will provide engineering design, catalysts, adsorbents, equipment, staff training and technical service for the project. Honeywell’s UOP technology to be used in this project includes: Honeywell’s UOP FCC process, to convert straightrun atmospheric gasoils, vacuum gasoils, certain atmospheric residues and heavy stocks recovered from other refinery operations into high-octane gasoline, propylene and light fuel oils; Honeywell’s UOP HF Alkylation process to produce a high-quality gasoline-blending component, typically referred to as alkylate; Honeywell’s UOP Caustic Merox process to remove sulfur from liquefied petro-

Construction leum gas (LPG) streams; Honeywell’s UOP Huels Selective Hydrogenation Process (SHP) to minimize acid consumption in the alkylation unit, produce 2-butene and maximize alkylate yields; and Honeywell’s UOP Butamer process to convert butane to isobutane, a primary feedstock used to produce alkylate in the HF Alkylation process. LANXESS has chosen Burckhardt Compression to deliver one process gas compressor for its chemical production site in Leverkusen, Germany. The compressor will be used to compress ethylene from 17 bara to 495 bara. In addition, Burckhardt Compression (Deutschland) GmbH has been awarded an order from LANXESS to revamp two existing process gas compressors. A gas-to-liquids (GTL) project in Uzbekistan was named OLTIN YO’L GTL at a formal ceremony in Tashkent involving representatives from the three jointventure ( JV) companies: Uzbekneftegaz, Sasol and PETRONAS. The naming of the JV project followed a ceremony to mark the start of infrastructure development by the Government of Uzbekistan at the proposed GTL plant site at Shurtan in the south of Uzbekistan. It also aimed at supporting the project schedule to enable a final investment decision, which is expected during the second half of 2013. When commissioned, the 38,000-bpd plant will produce a combination of GTL diesel and GTL naphtha and, in an important development in the application of GTL fuels, GTL kerosine for the aviation sector. Neste Oil has completed the first phase of its project to build a pilot plant for producing microbial oil. Plant construction is on schedule and on budget. The first phase will enable the growth of oil-producing microorganisms, and the following phases will concentrate on raw material pretreatment and oil recovery. The goal is to develop the technology so that it is capable of yielding commercial volumes of microbial oil for use as a feedstock for NExBTL renewable diesel. Commercial-scale production is expected by 2015 at the earliest. The pilot plant is expected to be fully complete in the second half of 2012, and

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Construction it represents an investment of approximately €8 million by Neste Oil. Microbial oil technology is attractive because of its efficiency and sustainability. Neste Oil has carried out pioneering research in the field and has applied for numerous patents covering its microbial oil technology. A number of partners have been involved in this work, including Aalto University.

pacity of several thousand metric tpy, is scheduled for completion in mid-2013. The investment, in the company’s largest production site worldwide, is in the mid double-digit million euro range. The new plant can make optimum use of the existing infrastructure and raw material supplies, and utilize synergies with the current polybutadiene plants in the Marl Chemical Park.

In May 2012, ThyssenKrupp Uhde won a front-end-engineering and design contract for a single-train polypropylene (PP) plant based on LyondellBasell’s Spheripol process technology for ZapSibNeftekhim L.L.C, a wholly owned subsidiary of SIBUR. The 500,000-tpy plant is planned to be constructed in Tobolsk, Russian. The plant will produce a wide range of highquality PP brands.

Middle East

Evonik Industries has laid the foundation stone for a new, large plant to produce functionalized polybutadienes in Marl, Germany. The plant, with a ca-

KAR Group has announced the third expansion of its Kalak refinery to 185,000 bpd. Process units and utilities are being provided by Ventech Engineers LLC of Pasadena, Texas. Ventech has provided modularized crude-distillation units, naphtha hydrotreaters, catalytic reformers, isomerization units and demercaptanization systems, as well as gas plants and supporting utilities to the project. This is the refinery’s third expansion, and is a continuation of Ventech’s modular construction methods. The first phase utilized 26 process modules to add 20,000 bpd of refining capacity to KAR’s

existing 20,000-bpd plant. The subsequent second phase provided an additional 60,000 bpd of total refining capacity and was completed in 2011. The third expansion consists of two 30,000-bpd modular complexes at the same site, as well as a 15,000-bpd condensate-processing facility. Once this latest phase is complete and operational, total capacity at the Kalak refinery will be over 185,000 bpd, and it will reportedly remain the country’s sole producer of unleaded gasoline, as well as the largest privately owned refinery in Iraq. Construction has started in Abu Dhabi, of what is said to be the largest Claus plant in the world. By the end of 2013, once construction is complete, Haldor Topsøe will supply the plant with 380 tons of the gas-treating catalyst TK-220. TK-220 is a CoMo hydrotreating catalyst specifically developed for treating tail gases derived from Claus or other similar units. The delivery is part of an agreement that Haldor Topsøe has made with BASF.

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Construction Asia-Pacific Jilin Qianyuan Energy Development has selected Chemtex, along with Black & Veatch’s patented PRICO LNG technology, to deliver a major liquefied natural gas (LNG) facility. Expected to be completed in late 2013, the 500,000-Nm3/d plant will reportedly be the largest of its kind in Northeast China. It will be delivered by the Chemtex/Black & Veatch team using a complete lump-sum engineering, procurement and construction package. The new facility will liquefy inlet pipeline natural gas. The LNG will be used primarily by trucks and other vehicles as an alternative fuel to diesel and petrol. In addition to utilizing PRICO technology, the plant integrates a nitrogen-stripping process. This will contend with high nitrogen levels in the pipeline feed gas. A special boil-off gas reliquefaction system will also be installed to prevent unnecessary fuel loss and increase plant efficiency. GTC Technology US, LLC, is licensing its GT-BTX extractive distillation

process to produce high-purity aromatics at Reliance Industries’ Jamnagar refining and petrochemical complex in Gujarat, India. This license is part of a consortium with CB&I Lummus Technology to supply technology for a multiunit complex for benzene and paraxylene production. Australia Pacific LNG has selected technology developed by Honeywell Process Solutions (HPS) Advanced Solutions business to deliver a data consolidation and reporting technology framework. The solution, built on Intuition Executive, will support the production, operations and asset-management functions of its world-class coal-seam gas (CSG) to liquefied natural gas (LNG) project. Origin, the upstream operator of the Australia Pacific LNG project, is the largest producer of CSG in Australia, supplying gas to power stations to produce electricity with lower greenhouse gas emissions. Australia Pacific LNG’s operational framework, built on Intuition Executive,

will support improved communication and data management, cross functional workflows and improved notification of key operational events to better analyze, prioritize, automate, and manage operational tasks and abnormal situations. Air Liquide has laid the first stone of a new hydrogen plant through a longterm agreement with Zhejiang Huafon Spandex Co., Ltd. (Huafon) to supply hydrogen for its 120,000-tpy cyclohexanone project located in Liaoyang Aromatics and Fine Chemical Park, Liaoyang city, China. Under the agreement, Air Liquide will invest in a new steam methane reformer (SMR) unit that will supply 13,000 Nm3/ hr of hydrogen, as well as steam to Huafon via pipelines. This new unit, which is expected to be commissioned by the end of 2013, uses Lurgi’s latest technologies providing high reliability, world-class safety and energy efficiency, and will be designed and manufactured by the Air Liquide engineering and construction team based in Shanghai.

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Hydrocarbon Processing | OCTOBER 2012 23

Construction SIBUR Petrochemical India, a subsidiary of SIBUR, has started operations in Mumbai, India. The company’s primary focus is to construct a butyl-rubber facility in Jamnagar. The plant, which will operate at a capacity of 100,000 tpy, is being built as a joint venture with Reliance Industries. The new subsidiary will work alongside Indian partners, as well as provide support to SIBUR employees coming to India to carry out installation and startup work at the new plant. Synthesis Energy Systems, Inc.’s syngas production facility for its Yima joint-venture ( JV) coal-to-methanol project in Henan Province, China, completed a test run of the first of three gasifiers currently under commissioning. During the commissioning phase, each of the three gasifiers will be operated under various test conditions to vet the gasifier and support systems to prepare the facility for commercial operation. This test was the first for this gasifier system operating on oxygen from the project’s air-separation unit.

Coal supplied by Yima was introduced into the gasifier, which operated for several hours while the Yima JV team successfully gathered data that will help in preparation for plant startup. The project will continue in the commissioning phase for several more weeks and is expected to move into commercial operation later in 2012. Intergraph has a contract with Santos for the use of SmartPlant Enterprise for Owner Operators (SPO), along with other SmartPlant Enterprise solutions. Santos will use Intergraph technology to manage its existing facilities and those in its Gladstone liquefied natural gas (GLNG) project in Queensland, Australia. Santos GLNG will use world-leading technology to process coal-seam gas (CSG) into LNG. The project is a partnership between Santos (operator), PETRONAS, Total and KOGAS. SPO enables Santos to address interoperability issues, while enhancing plant safety and reliability, quality and

productivity. In addition to SPO, Santos has implemented SmartPlant 3D, SmartPlant Foundation, SmartPlant Instrumentation, SmartPlant Electrical, SmartPlant P&ID and Leica Geosystems laser-scanning solutions, and also used the Intergraph Data Conversion Center for converting P&ID drawings into intelligent data. Besides the GLNG project, Santos will also use these design and engineering solutions for the entire life cycle of its facilities for both greenfield and brownfield projects, with plans to extend these solutions to its existing facilities within the Asia-Pacific region. CB&I has a contract with the Hebei Haiwei Group for the license and basicengineering design of a grassroots propane dehydrogenation unit to be located in Jingxian, Hebei Province, China. The unit will use the CATOFIN propane dehydrogenation process from Lummus Technology, which uses Süd-Chemie’s latest CATOFIN catalyst to produce 500,000 metric tpy of propylene. The unit is expected to start up in 2015.

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www.fourquest.com www FourQ FourQuest Energy focuses on servicing the needs of our clients and making a positive impact to thei their business. Motivated people are the most important resource of our company. import We take tak pride in providing safe, quality services to the en energy industry. A thor thorough understanding of our customers, their processes and issues, is the key to our accurate proces and th thoughtful approach. Having exceptional depth of experience and in-house technical expertise allows us to consistently meet or expert surpass client expectations. Find Us On: FourQuest Energy is an ISO 9001:2008 Certified and Registered Company

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CONSTRUCTION BOXSCORE UPDATE / ConstructionBoxscore.com COMPANY

CITY

PROJECT

EX CAPACITY UNIT

COST STATUS YR CMPL LICENSOR

ENGINEERING

CONSTRUCTOR

ASIA/PACIFIC China Kyrgyzstan Thailand India India Indonesia Kyrgyzstan

Jilin Qianyuan Energy Development Socar PTT Public Co Ltd Reliance Industries Ltd GAIL (India) Ltd BP Undisclosed

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LNG

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2013

Chuy Province Map Ta Phut Jamnagar Pata Papua Chuy Province, Kara-Balta

Refinery Waste heat recovery unit Refinery (1) Ethylene Dimerization LNG (3) Refinery

2 MMtpy None None 20 Mtpy 3.8 MMtpy None

P E P U U P

2013 2013 2014 2014 2017

GTC, Inc

OPTI Canada Inc Consumers Coop Refineries Consumers Coop Refineries

Fort McMurray, Long Lake Regina Regina

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2012 2012 2012

Shell | Fluor UOP UOP

Evonik Industries AG NKNK Sibur Sibur Sibur Khimprom Tobolsk-Polimer Haldor Topsøe Lukoil Lanxess

Marl Tatarstan Tobolsk Tobolsk Tobolsk Tobolsk Tyumen Nizhny Novgorod Leverkusen

Polybutadiene Ethylene Ethylene Polyethylene (4) Polypropylene Polypropylene Hydrogen Purifier EX FCC Gasoline Desulfurization Compressor RE

1 1.5 1.5 550 500 7 1.1

U F F F E F U U U

2013 2013 2013 2013 2013 2013 2013 2015

Paramaribo Cosoleacaque Barrancabermeja Tula

Refinery Cogeneration Hydrocracker Refinery

15 85 80 250

9000

U U E E

2013 2014 2016 2016

Jubail 2 Ind Zone Jubail Ind City Duqm Shiraz Ras Laffan

Complex Complex Petrochemical Complex Refinery Gas Processing

None 360 3 MMtpy 20000 230 bpd 1500 120 bpd 1300 None

E U P P E

2014 2016 2017

Ingleside

Ethylene Cracker

500 tpy

2016

EX BY

250 1000 7000

Chemtex Intl | Black & Veatch FW Fluor

Samsung Eng Axens

CANADA Alberta Saskatchewan Saskatchewan

EX EX

620 LTPD 22 Mbpd 30 Mbpd

1500

Fluor Mustang | IAG Mustang | IAG

EUROPE Germany Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Russian Federation Germany

Mt MMtpd MMtpy MMtpy Mtpy MMtpy MMtpy MMtpy None

INEOS INEOS LyondellBasell | INEOS UOP Axens | UOP

CB&I Linde Technip Linde | Vnipineft ThyssenKrupp | Linde-KCA Fluor | Linde-KCA Burckhardt Compression

LATIN AMERICA Surinam Mexico Colombia Mexico

Staatsolie Alpek Ecopetrol Pemex

EX RE

Mbpd MW Mbpd Mbpd

575

CB&I Lummus Axens

Aker Solutions Sener FW ICA Fluor

Saipem | Honeywell Sener Axens

MIDDLE EAST Saudi Arabia Saudi Arabia Oman Iran Qatar

Sadara Chemical Co. Sadara Chemical Co. OOC/IPIC NIORDC RasGas

Tecnimont Tecnimont

ABB | Linde Shaw

JGC

UNITED STATES Texas

Mexichem/Oxychem

The above projects represent only a fraction of data updated monthly in the Construction Boxscore Database. For more information please go to www.ConstructionBoxscore.com or contact Lee Nichols at 713-525-4626 or Lee.Nichols@GulfPub.com.

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Hydro, Inc. l HydroAire, Inc. l Hydro East, Inc. l Evans Hydro, Inc. l Hydro South, Inc. l HydroTex Golden Triangle l HydroTex Dynamics, Inc. HydroTex Deer Park, Inc. l CW Hydro, Inc. l Hydro Australia, Pty. Ltd. l Hydro Vietnam, Co. Ltd. l Safe-T Hydro, Inc. Hydro Scotford, Inc. l Hydro Middle East, Inc.

Reliability

HEINZ P. BLOCH, RELIABILITY/EQUIPMENT EDITOR Heinz.Bloch@HydrocarbonProcessing.com

Equipment life extension involves upgrades Every plant wants to get more service life from site machinery. Since about 1990, quite a number of start-up consulting companies have formed to advise clients on equipment life extension. These companies use different approaches: some apply large-scale, computer-based statistical methods, while others blend traditional estimates with risk-based analysis. All of these approaches have merit, but none of them can provide all of the answers with high precision. The key ingredients of any useful endeavor include reviewing the asset’s past failure history, examining nondestructive testing (NDT) data, and upgrading the weakest link. Failure history counts. Wherever failure history exists and

the failures’ root causes were analyzed, authoritative answers on remaining service life are possible. The same can be said for thoroughly evaluating NDT data, which can provide focus to determine remaining life. On stationary equipment and piping, wall thickness is of great importance. Loss of material decreases the allowable pressure rating. Corrosion and erosion can lower the safety of the equipment; thus, continued operation becomes risky. Thickness changes often occur at locations, such as elbows, where fluid flow changes direction. Changes in velocity such as at valves or near restrictions are of high interest. Some can be investigated with NDT methods, which certainly include X-ray imaging, among others. The extent of fluid-dependent corrosion can be estimated from coupons placed in piping and vessels. Pumps. For pumps, failure history and past repair data must be matched with a thorough understanding of upgrade measures that have been taken by successful “best-of-class” organizations. Advanced lube application will probably be part of it, as will the extension of oil replacement intervals now possible by synthetic lubricants and advanced bearing housing protection measures. To what extent superior bearings (ceramic hybrids) are of value must be determined on a pump-by-pump basis. Perhaps a set of angular bearings with unequal contact angles should be installed in your problem pumps. The symmetrical sets of angular contact bearings mentioned in the most widely used pump standard may not perform adequately. The extent that superior sealing technology (dual seals, as shown in FIG. 1) provides more value must be determined on a serviceby-service basis. As a general rule, the industry’s view about dual seals deserves to be reassessed. Sealing technology has made considerable progress in the past two decades. Virtually all present-day seals are cartridge-style configurations, and braided packing is being displaced by mechanical seals in the hydrocarbon processing industry (HPI), as well as in the power generation and mining industries.

FIG. 1. Dual mechanical seal in a slurry pump. The space between the sleeve and the inside diameter of the two sets of seal faces is filled with a pressurized barrier fluid—usually clean water. Source: AESSEAL Inc., Rockford, Tennessee, and Rotherham, UK.

However, not all manufacturers of mechanical seals use the same acceptance test procedure for their products. A widely applied industry standard stipulates using air as a test gas for mechanical seal tightness. Of course, these seals are ultimately intended for safe containment of flammable, toxic or otherwise hazardous liquids. While the standard’s expectation is that leakage from these seals does not exceed 5.6 gm/hr, recent tests showed that merely following this easy testing routine can actually allow orders of magnitude more liquid to escape. It is, therefore, advisable to question seal vendors on the matter and to purchase only products that meet the purchaser’s safety and reliability requirements. We all want seals to leak no more than 6.5 gm/gr when first installed on pumps. Hydrocarbon Processing | OCTOBER 2012 29

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Reliability

Lubrication systems. Lubricant application and standby bearing preservation are especially important in humid coastal climates and in dust-laden desert climates. Oil mist is the answer. The settling of foundations and pipe supports should be addressed. For steam turbines, the blade stresses and water quality must be compared with those units in successful longrunning installations elsewhere. In gearboxes, the remaining service life is largely examined

by tooth loading (stresses on tooth face) and temperature measurements. In all instances, synthetic oils from the most experienced oil formulators will greatly extend gear life. Oil additives are everything. They drive both cost and service life. Oil cleanliness is equally important. Certain warehouse spares (gears, electric motors, etc.) should be upgraded, if important. If doing so, it is likely to speed up re-commissioning after an unanticipated future shutdown. Compressors. For compressors, engineers should consider the mentioned points. Valve technology and piston velocity are important comparison-worthy parameters on reciprocating compressors. Onstream performance tracking and prior sealing technology are important for centrifugal compressors, etc. They determine seal system upgrade potential. Never overlook couplings and the work procedures used to attach couplings to shafts. They tell a lot about remaining run length. Consultants. Whether one ultimately receives life extension

guidance from individual consultants or from billion-dollar consulting giants with applicable experience is of no consequence, as long as there is the one common thread: Determining where upgrades are possible. Upgrades are critical to imparting longer life to existing equipment, and they can often be accomplished at relatively low cost. Assessments of remaining life should include detailed advice on how to upgrade weak links, which implies: • The expert authoritatively spells out recommended upgrade components • Recommended upgrade procedures are explained • Facilities recommended to do the upgrading are defined. In short, the entity involved in advising you on equipment life extension must understand the feasibility of component upgrades. Component upgrading is one of the keys to life extension and deliverables that should be contractually agreed upon with the upgrade provider. Be sure that the consulting company you’ve asked to give advice on equipment life extension includes these deliverables.

With over 50 independent subsidiaries and more than 220 engineering and sales ofﬁces spread across the world, SAMSON ensures the safety and environmental compatibility of your plants on any continent. To offer the full range of high-quality control equipment used in industrial processes, SAMSON has brought together highly specialized companies to form the SAMSON GROUP.

HEINZ P. BLOCH resides in Westminster, Colorado. His professional career commenced in 1962 and included long-term assignments as Exxon Chemical’s regional machinery specialist for the US. He has authored over 520 publications, among them 18 comprehensive books on practical machinery management, failure analysis, failure avoidance, compressors, steam turbines, pumps, oil-mist lubrication and practical lubrication for industry. Mr. Bloch holds BS and MS degrees in mechanical engineering. He is an ASME Life Fellow and maintains registration as a Professional Engineer in New Jersey and Texas. A01120EN

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SAMSON AG · MESS- UND REGELTECHNIK Weismüllerstraße 3 60314 Frankfurt am Main · Germany Phone: +49 69 4009-0 · Fax: +49 69 4009-1507 E-mail: samson@samson.de · www.samson.de SAMSON GROUP · www.samsongroup.net Hydrocarbon Processing | MONTH 2012 31

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Integration Strategies

BARRY YOUNG, CONTRIBUTING EDITOR Byoung@ARCweb.com

Recent trends shape the future of DCS Several trends have already impacted the distributed control system (DCS) market and are likely to continue over the next few years. These include both product- and technologyrelated trends and general industry trends. More intelligent I/O. The DCS input/output (I/O) subsys-

tem is responsible for inputting hundreds or often thousands of different process measurements and other inputs into the system, and outputting control signals to a large number of valves, actuators, motors and other plant final-control elements. I/O represents one of the most significant parts of the DCS. Traditionally, I/O is a significant cost element. However, DCS suppliers are working to reduce both the cost and complexity of their I/O systems by incorporating more intelligence and programmability into the devices. Shift in I/O type. Fifteen years ago, the traditional process

analog input came from a sensor producing a 4-mA to 20-mA analog signal, and the typical analog output was a 4-mA to 20-mA signal. Discrete signals involved various combinations of voltages and currents. Each signal type had a dedicated type of circuit board for the individual signals. Today, in a greenfield plant, most of the I/O supplied is on some type of bus network. Brownfield plants are also installing more bus I/O. However, with the large installed base of traditional 4-mA to 20-mA I/O, the transition is very slow. Major expansions or revamps in brownfield plants consider bus I/O when the sensors and final control elements are also part of the project. There is also a growing trend to adding more wireless I/O and associated field devices, particularly for process and equipment monitoring applications. Need for network consulting services. As the lines be-

tween automation and IT are blurring with increasing usage of commercial off-the-shelf (COTS) technology, the network infrastructure of the DCS and the network architecture for plant information are becoming increasingly more intertwined. End users now often rely on the expertise of suppliers for consulting to set up these networks in a safe and secure manner. Virtualization. DCS suppliers started incorporated server vir-

tualization a few years ago. Common uses of this technology include engineering development and simulation for training. Virtualization is not appropriate for all parts of the DCS. Sometimes, the dedicated hardware will perform a given task better than a virtual server. A good example is the real-time process controller in a DCS, where speed, determinism and high reliability are major design considerations for the operation and safety of the plant. Conversely, a virtual server performing many

applications on one box can be a good choice for offline applications such as control configuration, simulation and training. Cyber security. With more open and interoperable, largely COTS-based automation systems, cyber security is becoming more important as end users struggle with potential risks, both internal and external to the DCS. Most suppliers now address this threat with active programs, either inhouse or through partnerships. As part of a “defense-in-depth” cyber-security strategy, network fire walls and strategically placed switches are required to help prevent the propagation of external viruses and intrusions. Internal threats from disgruntled employees or other internal access points must be addressed with such things as USB locks and software to monitor internal automation system network activity. Furthermore, network maintenance practices that are common in the IT world—such as automatic software updates, anti-virus updates and bug fixes—must be modified for the mission-critical, 24/7 industrial environments in which DCSs typically operate. Mobility. Just as people today find it hard to live without their smartphones in their daily lives, increasingly, process operators and production supervisors are relying on the ability to “access data anywhere, anytime” to do their job functions. DCS suppliers address this trend by supplying tablet technology for roving operators and smartphones for alerts and condition monitoring. This trend toward increasing mobility will grow in importance in the future. Cloud computing. There has been much talk in the industry

about developments underway to move selected DCS applications “to the cloud,” a reference to moving applications to remote, Internet (public) or intranet (private) servers. However, the control-automation industry is very conservative by nature, and, presently, this trend is just talk. ARC believes that, ultimately, selected DCS applications will migrate to private and, in some cases, even public “clouds.” For now, end users are wary. BARRY YOUNG has over 25 years of professional experience in the process control and industrial automation industry. Prior to joining ARC, he served as a project manager for New England Controls, where he helped design and implement automation solutions for a variety of high-profile clients in the life sciences, utility, pulp and paper and other industries. Prior to New England Controls, he handled a variety of responsibilities within the global Invensys/Foxboro Company organization. Mr. Young has a BS degree in management engineering from Worcester Polytechnic Institute, and has completed MBA courses at Bryant University. He is a member of the Project Management Institute. Hydrocarbon Processing | OCTOBER 2012 33

| Special Report PROCESS CONTROL AND INFORMATION SYSTEMS For as long as hydrocarbon processing industry (HPI) facilities have been processing crude oil, natural gas and intermediates, there have been instruments in place to assist plant operators in measuring, recording and controlling pressures, flows, levels, temperatures and other process variables. Much has changed since the early days of the HPI. With the application of digital control and computer-integrated manufacturing, facilities have been able to automate control of processing units, directly analyze product streams and initiate action from a central control room. Since the 1980s, technology developments in field instrumentation design, software, simulation models and more have provided even greater opportunities to further implement process monitoring and control and, ultimately, optimize plant operations. The goal for any process control project is to increase profitability; to this end, plant optimization and process automation are discussed in this monthâ&#x20AC;&#x2122;s Special Report. Photo courtesy of ABB Process Automation.

Special Report

Process Control and Information Systems A. G. KERN, Consultant, Houston, Texas

Update hydrocracking reactor controls for improved reliability At present, hydrocarbon pricing pays strong dividends for hydrocracking (HC), which leverages low-cost hydrogen (H2 ), sourced from today’s abundant natural gas supply, into high-value liquid fuels. Among hydroprocessing (HP) units, HC units have the greatest H2 uptake, with a typical liquid volume swell of 10%. This earns in the range of $100 million (MM) per year for a 30,000-barrel-per-day (bpd) HC unit based on volume swell alone, in addition to the usual value upgrade of HC conversion. As a result, hydrocrackers—which are already one of the hydrocarbon processing industry’s most demanding process control challenges—are being pushed to greater limits. HC reactors operate at elevated temperatures and pressures, making safety a constant concern. Recovering from temperature upsets can take hours, and recovering from a complete depressurization takes days. Reactions are exothermic, meaning that even minor disturbances in feed, heater or quench controls can rapidly escalate to an urgent situation. For these reasons, HC controls have always demanded vigilant attention to design detail, management of change, and operability. Any oversights can result in, or fail to prevent, depressurization events. Key improvements, on the other hand, can bring large gains in refinery profitability and reliability. For the past two decades, hydrocracker process control strategy has focused on installing automatic depressurization systems and multivariable predictive control (MPC). While this equipment has brought important gains, experience shows that it leaves many gaps in excursion control and depressure prevention. This article presents an updated hydrocracker control model that robustly addresses traditional hydrocracker control challenges, overcomes outdated hydrocracker control paradigms, and allows hydrocrackers to operate safely, reliably and profitably under today’s demanding conditions.

desired amount of cracking (or “conversion”), which is borne out in the downstream fractionation section product spreads. Maximizing conversion means operating at one or more of the quench constraints. These include a maximum quench valve position, chosen to ensure ample reserve quench should an exothermic excursion occur, and a maximum bed temperature rise, which indicates high cracking severity and increased risk of a rapid onset excursion. Recent price trends in crude oil and natural gas have shifted HP economics. The price of natural gas has declined, while the price of crude oil and liquid fuels has greatly increased. H2 consumed through a hydroprocessing complex swells the liquid yield, effectively converting a low-cost feedstock into a highvalue product. Among HP units, HC has the greatest H2 uptake, typically around 1,700 standard cubic feet (scf) of H2 per barrel (bbl), with a resulting liquid volume swell of 10%. Therefore, the gross profit margin from volume swell alone is in the range of $100 million (MM) per year for a 30,000-bpd HC unit. This is based on H2 sourced from natural gas at $4/thousand scf, and valuing product fuels at $120/bbl. Past strategies and present gaps. Over the past 20 years,

hydrocracker process control strategy has focused on installing Combined feed (H2 and oil) from heaters

Bed 1 quench

Manual HC DEPR depressure

Distribution TC IN-1A

TC IN-1

Heater control

Catalyst Bed 1 Contact and distribution

TC IN-2

HC process and economics. FIG. 1 shows a common hydro-

cracker configuration. Heated oil and excess H2 enter a vertical downflow reactor with multiple fixed catalyst beds. The catalyst promotes cracking and hydrogenation of larger hydrocarbon molecules, such as gasoils, cycle oils and coker oils, into lighter, more valuable molecules, such as diesel, jet fuel and naphtha. The overall reaction is exothermic, so temperature increases as flow passes through the bed. Between beds, cold H2 quench gas is introduced to cool the reaction mix. In this way, the reactor is a succession of cracking beds followed by quenching (FIG. 1 depicts three beds, but often there are more). The overall objective is to achieve the

Flare

RO Recycle H2 Separator PC SEP

Catalyst Bed 2

LC SEP

Contact and distribution TC IN-n

Quench H2

Fractionation

Catalyst Bed n Support

Feed/effluent exchangers

FIG. 1. Simplified HC reactor piping and instrumentation diagram with “as-purchased” controls. Hydrocarbon Processing | OCTOBER 2012 35

Process Control and Information Systems automatic depressurization systems and MPC. While these are important, experience now shows that they leave many gaps in excursion control and depressure prevention. Auto-depressure controls serve to vent reactor systems to flare in the event of an uncontrolled exothermic excursion, to halt the reaction and prevent vessel temperatures from exceeding metallurgical limits. Temperatures during an exothermic excursion sometimes increase tens of degrees in as many seconds. Industry history shows that manual depressure systems are often not used according to written procedures and

Auto-quench is a DCS control, but it can be one of the most important functions in a refinery, since it is the final layer of depressure prevention. that personnel at all levels—managerial, supervisory and operational—can have difficulty balancing production goals with safe use of manual depressure systems.1 This illustrates why auto-depressure—made possible by more reliable thermocouple-based temperature-measurement systems and improved algorithms for excursion detection and temperature-measurement quality handling—has become essential. The difficulty with auto-depressure controls is that they tend to act much sooner than traditional, manually initiated systems, and depressuring is to be avoided whenever possible, except as a final layer of safety. Depressuring a reactor brings the unwelcome prospects of prolonged restart, impact on other refinery units, large economic losses, thermal and mechanical stresses to the reactor and associated equipment, environmental flaring violations, and potential harm to the company image in the community and in the industry. The necessary message that often fails to accompany auto-depressure projects is the need for better excursion control to avoid reaching auto-depressure conditions in the first place. MPC technology has brought improvements in reactor bed temperature balancing, weighted average bed temperature (WABT) control, and coordination between reactor and fractionator sections (conversion control). However, MPC lacks the speed, reliability and control features necessary to adequately respond to most hydrocracker disturbances before they result in an excursion, or to contain an excursion before it leads to depressure, or to do so in a manner that minimizes overall impact on reactor temperatures and resulting lost production. 1. Auto-depressure 2. Auto-quench 3. Excursion control

ent

Safe

4. Bed outlet control 5. WABT and rate control

rsio

Excu

ainm cont

ion erat p o al

Norm

mi Opti

6. Conversion control FIG. 2. HC reactor control model showing “layers of control.”

36 OCTOBER 2012 | HydrocarbonProcessing.com

zati

Excursion scenarios. Hydrocracker operators are always aware of the many potential excursion initiators. On the other hand, when designing controls (and even during hazard analysis), there is a common tendency to downplay the likelihood, severity and actual history of many excursion scenarios.1 For a hydrocracker (or for any critical process control), an effective approach is for a multidisciplinary team to consider each scenario and the most appropriate control system response. Common excursion scenarios include: • Loss of oil feed. A feed pump trip normally results in a strong excursion unless quickly quenched, because the oil stops moving through the bed and instead “cracks in place,” never reaching the quench zone. • Heater operations. Adding burners, fuel gas upsets, draft or oxygen upsets, etc., can cause feed temperature spikes, triggering an excursion. • Maldistribution. Bed inlet maldistribution can cause erratic quench controller behavior, especially if a single measurement point is used for control or if inter-bed redistribution internals are not functioning properly. Maldistribution of flow through the catalyst bed creates localized low flow conditions and “hot spots” where excursions can take hold. • Production changes. Although operating procedures are designed to implement changes conservatively and safely, excursions commonly occur during changes to feed rates, feed type or temperature (i.e., conversion). Additional potential excursion triggers are listed in Reference 1. Complex refineries with a variety of feed and product types can be subject to these hazards on an essentially continuous basis. Understanding these causes helps build better controls; however, the control design must also provide effective excursion control, regardless of the cause. Layers of control. FIG. 2 is a hydrocracker reactor control model that addresses safety, depressure prevention, and excursion control, along with normal operating objectives and optimization. The overlapping of layers indicates robust reliability. For example, excursions may be contained and controlled by Layer 4, 3 or 2 before ever reaching Layer 1 (depressure). In addition, Layers 4 and below are implemented in the baselayer control system, thereby maximizing responsiveness, reliability and operability. • Auto-depressure on high temperature is becoming established as an industry best practice. Key design decisions include whether to implement auto-depressure in the safety instrumented system (SIS) or the distributed control system (DCS); whether to depressure on high temperature rate-of-change (in addition to high absolute temperature); and how to robustly handle low-quality temperature singularities among the bed outlet thermocouple arrays to avoid unnecessary or nuisance depressurization events. • Auto-quench causes the quench valves to open on high excursion temperature to avoid reaching the depressure limit. It may also trigger preemptively on feed pump trip and initiate heater minimum fire logic. Auto-quench design is a balancing act: it should be robust, like a safety function, but without being so heavy-handed as to result in an extended recovery time; it must trigger early enough to avoid reaching depressure, but

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without triggering unnecessarily; and it should not interfere with, or be defeated by, quench controls in manual mode. Although auto-quench is a DCS control, conceptually it can be one of the most important functions in a refinery, since it is the final layer of depressure prevention. • The excursion control layer is designed to handle excursions as routine control disturbances, when possible, to minimize their impacts. This renders most excursions as non-events, such that they go largely unnoticed, except perhaps by the DCS operators. In the past, operators remained alert to take manual control in the event of an excursion, but with excursion controls, operators learn to keep them in the correct mode to ensure reliable automatic response. As one operator noted, “These are controllers that work for us, and not the other way around.” The excursion control layer comprises a number of traditional advanced regulatory control (ARC) techniques applied to the bed inlet, bed outlet and heater controllers. An important aspect is converting the Bed 1 quench valve (“TC-IN-1A” in FIG. 1) to a bed outlet temperature controller (“TC-OUT-1A” in FIG. 3) and coordinating its action with heater control. This critical valve is often configured problematically (as in FIG. 1), so that it does not respond to an excursion and, when used, can cause the heater(s) to counter with increased firing. • In retrospect, operating an exothermic reactor without bed outlet control defies common sense, although it is a common practice, especially when MPC is switched off, detuned, clamped or over-constrained. Even if the excursion control features are absent, outlet control at least helps prevent many gradual process variations from reaching excursion thresholds. It also brings increased stability to bed outlet temperatures, WABT and conversion. However, simple outlet control—sans

Combined feed (H2 and oil) from heaters Heater control Bed 1 quench WABT and Conversion rate control control AC CONV

TC WABT

Model- Custom based or modelbased

TC AQ

Autoquench Quench H2

TC out-1A

TC IN-1

TC out-1 TC IN-2 TC out-1 TC IN-n TC out-n

Flare Auto/manual depressure Distribution Catalyst Bed 1 Contact and distribution

HC DEPR

RO Recycle H2

Separator PC SEP

Catalyst Bed 2

LC SEP

Contact and distribution

Fractionation

Catalyst Bed n Support

Feed/effluent exchangers

FIG. 3. HC reactor piping and instrumentation diagram with upgraded controls.

the excursion control features, including MPC—usually will not contain an excursion once it begins. • MPC-based WABT control, when implemented in the new model (FIG. 3), would write to the bed outlet controller setpoints rather than to the bed inlets, as is traditional. The bed outlet controllers provide base-layer stability and excursion control, while MPC provides traditional WABT control and constraint management. As an alternative, WABT control can be implemented as a custom algorithm, providing greater flexibility in how the constraints are managed. This also facilitates a Hydrocarbon Processing | OCTOBER 2012 37

Process Control and Information Systems metric, frequent, minor excursions indicate the increased likelihood of a full-blown excursion and potential depressure event. A graph like that shown in FIG. 4 is a good candidate for visibility on a large control-room screen, as a means of sustaining improvement and awareness. Recommendations. A guiding tenet in the evolution of these

FIG. 4. Improvement in excursion control.

variable WABT ramp rate that can be both faster and safer during startup and recovery operations, capturing extra hours of on-target production. Since the base-layer controls handle stability, this custom WABT algorithm is similar to, and no more complex than, a traditional heater pass balancing control. • Conversion control involves moving the WABT setpoints based on fractionation section product spreads. MPC is a good choice for handling the long response times involved, although a rudimentary custom algorithm can also be used. A limitation is that feed quality changes are the primary disturbance, and they are often much “bigger than the handles,” since WABTs can only be moved gradually and within limits. Another key consideration is a smart-conversion calculation and scheduler, so that fractionator disturbances are not back-propagated to the reactor section. Metrics. Excursions are commonly quantified as the difference between real-time reactor bed temperatures and recent (heavily filtered) values. The excursion value reflects any short-term temperature rise; i.e., the severity of an excursion. At steadystate, this value will be zero, and at operating conditions (if procedures are carefully followed and no excursions occur), it will always be less than the prescribed maximum hourly rate of change; e.g., less than around 5°F. Since modern hydrocracker reactors may have a dozen or more thermocouples per bed, a common practice is to calculate the highest temperature of each bed for monitoring and alarming, for ease of operation, and to avoid alarm floods when excursions occur. FIG. 4 is an example of the long-term trend of the highest excursion temperatures for each bed of a two-stage reactor. The vertical axis shows excursion severity (for example, increments of 5°F). Excursions below a severity of 1 reflect routine daily operation. Excursions with a severity between 1 and 2 may occur daily, weekly or monthly, depending on the quality of operation. Excursion controls should take effect at this level. Excursions greater than a severity of 2 are increasingly serious and, in many cases, warrant near-miss investigations to prevent recurrence. These investigations often lead to the types of control improvements described here. FIG. 4 provides a meaningful metric of progress and of ongoing quality of operation. As controls are upgraded, the frequency and severity of excursions decreases. As with any safety 38 OCTOBER 2012 | HydrocarbonProcessing.com

controls was to utilize quench, heater and other controls as advantageously as possible under all circumstances, to contain excursions and avoid reaching auto-depressure conditions. This led to many creative and sensible ideas. The main challenge was not in the difficulty of designing new controls reflecting these ideas, but in overcoming entrenched paradigms about the old controls, even though they were outdated or not sensibly configured in many cases, such as the conflict between the Bed 1 quench and heater controls, and the lack of reliable bed outlet control. Control layers 2 through 4 were implemented with standard DCS functionality, bringing cost and engineering advantages. Another practical benefit is operability, since these controls present to the DCS console operator and behave as conventional cascaded controls, requiring minimal new concepts and training. MPC is often considered a comprehensive solution for the types of control concerns raised here; however, none of the critical excursion control, depressure prevention or auto-quench functions are of the type provided by MPC. For design or hazard and operability study (HAZOP) purposes, it is usually better to view MPC as a gradual constraint pusher, rather than as a reliable disturbance handler. This distinction is important on any process, but especially for hydrocrackers, where a robust response can make the difference between an online reactor and a depressured reactor, in a matter of minutes. Process control could benefit by borrowing from safety system practice and convening a multidisciplinary team to review critical process upset scenarios and arrive at the most appropriate and advantageous automatic control response. While many processes do not have the rapid downside potential of HC reactors, the general principles of maximizing on-target production and avoiding safety function thresholds under upset conditions make this approach a good practice for any refinery unit. The traditional practice of operating high-pressure, hightemperature, exothermic reactors without reliable, nonlinear bed outlet temperature control is a paradigm that industry should proactively remedy. The auto-quench, excursion control and bed outlet layers should join auto-depressure as industry best practice for all hydrocrackers. 1

LITERATURE CITED EPA Chemical Accident Investigation Report, Tosco Avon Refinery, Martinez, California, November 1998.

ALLAN KERN has over 35 years of process control experience and has authored dozens of papers on multivariable control, inferential control, safety systems and distillation control, with a focus on practical process control solutions and effectiveness. He is a professional control systems and chemical engineer, a senior member of ISA and a graduate of the University of Wyoming. Mr. Kern is a consultant, and he can be contacted at Allan.Kern@APCperformance.com.

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Special Report

Process Control and Information Systems J.-M. BADER and G. ROLLAND, Axens, Rueil-Malmaison, France

Use a systematized approach of good practices in pygas hydrogenation via APC As illustrated in FIG. 1, steam crackers produce many basic building blocks for the polymer industry, along with aromatic-rich gasoline (pygas). When naphtha is used as the feed for cracking furnaces, pygas yields increase significantly. TABLE 1 shows the typical pygas yield and composition for naphtha cracking. Pygas is a large contributor to benzene production capacities. Before the pygas can be routed to downstream units (aromatics extraction, etc.), unstable compounds such as diolefins and styrenics must be removed. Also, olefins and sulfur must be eliminated to ensure that final products will meet their specifications. This pygas treatment is achieved through hydrogenation steps. However, if the pygas treating is not optimized, then other undesired processing operations—including hydrogen flaring, reactor channeling, poor use of the second beds, and other issues—have a greater possibility of occurring. Optimizing control. Advanced process control (APC) offers a solution to systematize implementing best practices, avoid mis-operation, and generate substantial benefits. The following case study describes the operational improvements and steps necessary when applying APC. This case study will apply actual plant data to demonstrate and quantify the benefits attainable from APC installations.

PROCESS A case study will describe the industrial results obtained with the two-stage pygas hydrogenation process (PGH), as TABLE 1. Typical naphtha cracker pygas (C5–200°C) yield and composition

shown in FIG. 2. This PGH unit includes a first-stage process (GHU-1) to improve the stability of the raw pyrolysis gasoline by selectively hydrogenating diolefins and alkenyl compounds, thus making it suitable for further processing in the second stage. The reaction is carried out mainly in the liquid phase on a specific catalyst in a fixed-bed reactor. The selected operating conditions maximize conversion of the diolefins and alkenyl aromatics, while minimizing the formation of heavy products by polymerization. These operating conditions minimize aromatics loss. In the second stage of PGH (GHU-2), the C6–C8 heart cut is further processed to prepare a feedstock suitable for H2 + CO C2 C3 Steam cracker

Feed

C5 – C10

Propylene

C3 hydrogenation

Butene, butadiene

C4 hydrogenation

C5 – BTX –C9+

Pygas hydrogenation

Fuel

FIG. 1. Pygas production and other products from a steam cracker. H2

Pygas feed D

Rx 1

Paraffins + naphthenes

11.8

Oleﬁns

5.5

Dioleﬁns

18.1

Benzene

Toluene

13.9

Xylenes

7.2

Styrene

Total aromatics

Ethylene

C2 hydrogenation

200° +

Composition, wt%

C9+ aromatics

High-purity hydrogen

Methanation

C5 Q

C9+

Rx 2

C 6 – C8 A

12.5 64.6

FIG. 2. Pygas hydrogenation flow scheme. Hydrocarbon Processing | OCTOBER 2012 41

Process Control and Information Systems aromatics recovery, by selectively hydrogenating the olefins and removing sulfur via hydrodesulfurization (HDS). The reactions are conducted through a series of specific catalysts in fixed-bed reactors. The operating conditions are selected to prevent aromatics losses by hydrogenation and to minimize heavy product generated by polymerization.

HYDROGEN NETWORK For APC of the hydrogenation processes, the hydrogen network plays a major role, and it needs to be studied carefully. Both pygas-treating stages consume hydrogen. Several configurations are possible to supply hydrogen to the pygas reactors, as shown in FIG. 3: • High-purity hydrogen option • Low-purity hydrogen option • First-stage purge in the second stage. One concern that cannot be ignored is that other facility processes are also hydrogen users, such as selective hydrogenation of C2, C3 and C4 streams. In a situation of low-hydrogen availability, these processes have priority. As a consequence, pygas can reach a situation in which the first stage is temporarily operated with insufficient hydrogen, which has negative consequences on process performance and catalyst life. When excess hydrogen is available, it is important to reduce wasteful To ﬂare

hydrogen flaring and to improve pygas operation by utilizing all available hydrogen. Efficient pygas operation can ensure that the best use is made of available hydrogen.

POSSIBLE OPERATING IMPROVEMENTS There are four key areas that have the potential to deliver operational improvements. These areas are: improving first-stage product quality, reducing the risk of channeling, maximizing second-bed usage, and optimizing global hydrogen usage. Improving first-stage product quality. The product enter-

ing the second stage must be hydrogenated to the correct level to prevent polymerization of any remaining diolefins or alkenyl aromatics in the second-stage reactor, which is operated at a higher temperature and in the vapor phase. If the hydrogenation process lacks sufficient hydrogen or has a low temperature profile, there will be a high tendency to form gums at the inlet of the second-stage reactor, thus generating unacceptable pressure drop and performance reduction. A good indicator of the hydrogenation of diolefins or alkenyl aromatics is the styrene content of the first-stage reactor product. Normally, the styrene specification is set at 1,500 ppm to efficiently protect the second-stage catalyst. In addition, reasonable catalyst cycles are followed. FIG. 4 presents a statistical distribution of the styrene content at the outlet of the first-stage reactor without APC. The

C2 hydrogenation High-priority H2 consumer

FC Pygas feed

C3 hydrogenation

Liquid load Diluent

C4 hydrogenation FC = Manipulated variable Liquid load = Controlled variable

H2 network

FC Quench

Rx 1 Product to 2nd stage

H2 purge Pygas feed

1st stage

To BTX extraction

2nd stage

FIG. 5. Diluent flow adjustment in first-stage reactor.

FIG. 3. Hydrogen network summary.

Feed ﬂow

Temperature profile in first bed of first-stage reactor

Reducing the channelling On speciﬁcation

No APC APC on

Off specs

Large giveaway 500 600 700 800 800 1,000 1,100

1,300

1,500

1,700 1,800 1,900 2,000

Styrene distribution, ppmwt

FIG. 4. Typical styrene statistical distribution in a first-stage reactor outlet (ppmwt) from an online analyzer.

42 OCTOBER 2012 | HydrocarbonProcessing.com

80 85 90 95 First bed temperature of ﬁrst-stage reactor, °C

FIG. 6. Improving from bumpy to smooth temperature profile with APC.

100

Process Control and Information Systems histogram is characterized by a small number of off-specification values that could have damaged the second-stage catalyst. Also, a large proportion of product was well below the required specification. This over-quality is translated as a cost or giveaway due to the unnecessarily high reactor temperature in the first stage that would reduce the catalyst cycle. Reducing channeling by appropriate diluent flow. In the first-stage reactor, the flow entering the reactor is mainly in liquid phase, and it is constituted of fresh pygas feed, diluent cooled and recycled from the first-stage reactor outlet and hydrogen makeup, as shown in FIG. 5. If the diluent flow is too low, then the hydraulic loading of the catalytic bed may become too small, and thus possibly cause channeling. If the diluent flow is too high, then the velocity in the reactor may be excessive. More importantly, it will lower the average bed temperature, thus a higher inlet temperature will be required to maintain performance. The higher operating temperature will negatively impact the stability of the catalyst. The total flow to the first-stage reactor (fresh feed + diluent), also called â&#x20AC;&#x153;liquid load,â&#x20AC;? must be adjusted to an optimal target value close to the design value to produce the most continuous temperature profile. This condition is illustrated in FIG. 6. With an inappropriate liquid load, the irregular temperature profile reveals the occurrence of channeling. Maximizing second-bed usage by quench flow. In both pygas reactors (first and second stage), there is usually a quench injection between the first and second beds, to control the reactor temperature profile. The quench flow is often kept too high by panel operators to prevent temperature runaways. In the first-stage reactor, as illustrated in FIG. 7, the consequences of excessive quench flow are a lower bed Î&#x201D;T, which results in lower hydrogenation levels in the second bed. This leads to increased styrene content in the product. To compensate for this case, the first-bed temperature is frequently increased, which is detrimental to catalyst life cycles. APC objectives for the first-stage reactor include the balance of the hydrogenation between the first and second reactor beds.

Optimizing hydrogen usage by minimizing flaring. This principle is described in FIG. 8. The hydrogen-network purge to the flare is piloted by a hydrogen-network pressure controller. If the valve of this pressure controller is not fully closed, then hydrogen is wasted and sent to flare. In this case, it is possible to increase hydrogen flow to the pygas unit, until the pressure controller valve is fully closed, thus optimizing usage of all available hydrogen In reality, the hydrogen-network pressure-control strategy implemented in the distributed control system (DCS) can be much more complex than presented in FIG. 8. Using in-depth knowledge of DCS capabilities, a new method was developed to minimize hydrogen loss and further increase the hydrogen makeup for the pygas unit without affecting the network pressure. This approach uses pressure controller parameters (setpoint, process value, valve opening, etc.) and dynamic models derived from step-test data. The benefits from this approach are to deliver more hydrogen to the pygas unit and thus improve hydrogenation performance.

APC STRATEGY This novel approach on APC strategy was successfully applied to optimize industrial pygas process operations. It incorporated several key control methods to improve the hydrogenation process: To ďŹ&#x201A;are H2 available

PC C2 hydrogenation

Send all ďŹ&#x201A;ared H2 to pygas

C3 hydrogenation C4 hydrogenation

H2 network

H2 purge Pygas feed

1st stage

To BTX extraction

2nd stage

FC = Manipulated variable H2 available = Infered controlled variable

FIG. 8. Using hydrogen network information to maximize hydrogen supply to pygas.

Styrene in product Diluent

Quench

Diluent

1st stage

2nd stage

1st stage reactor 2nd bed Î&#x201D;T

Lab data Ä&#x2018;ĆŤ 05.!*! Ä&#x2018;ĆŤ . Ä&#x2018;ĆŤ !*/%05 Feed quality estimation

Quench ďŹ&#x201A;ow steps

FIG. 7. Effect of quench flow changes (during 14 hours) in the second bed of first-stage reactor.

Quench

Lab data

Reactor model

05.!*! dioleďŹ ns .

Reactor model

BI index aromatics

Periodical manual catalyst activity update

FIG. 9. Inferential model to maximize hydrogen management while minimizing styrene content. Hydrocarbon Processingâ&#x20AC;&#x201A;|â&#x20AC;&#x201A;OCTOBER 2012 43

Process Control and Information Systems Maximize feed. Common practice is to place an intermediate product tank between the steam cracker and the pygas unit. The volume of this tank can usually absorb one day’s production of pygas. The inventory of this tank should be minimized to reduce the risk of polymerization of unsaturated components present in the raw pygas, until downstream constraints have been saturated. Use all available hydrogen. Optimize the global hydrogen management. First-stage reactor. The first target is to do ultra-deep hydrogenation of diolefins and alkenyl aromatics, by controlling the styrene content, as measured by an online analyzer, in the first-stage product. The next step is to stabilize reactor operation by controlling the reactor liquid load at an ultimate level to avoid channeling. Hydrogen partial pressure is maximized to promote hydrogenation by increasing the reactor pressure while maximizing the dissolved hydrogen fraction in the liquid phase. The process is operated to ensure a minimal gas purge flow to prevent concentration of inert species in the hydrogen recycle gas. (This is applicable if the unit is equipped with a recycle-gas compressor.) The temperature profile is optimized, using reactor inlet temperature, diluent and quench flow to prevent temperature runaway, to balance reactor ΔT between the two beds, and to maximize catalyst cycle length.

Styrene online analyzer

Fractionation. APC needs to identify the right compromise

between the quality of the separation and energy savings. Second-stage reactor. The first target is to perform com-

plete hydrogenation of olefins and sulfur removal by controlling the bromine index (BI) and sulfur content of the reactor effluent. The next step is to minimize hydrogenation of aromatics by avoiding unnecessarily high-temperature process conditions. Finally, stable reactor operation will be achieved by the control of reactor ΔT and hydrogen-recycle gas density.

PYGAS INFERENCE AND OPTIMIZER The pygas inferential model proposed for APC, as shown in FIG. 9, is based on highly evolved kinetic models that enable online styrene content and BI estimation, and, consequently, reactor optimization. Laboratory analyses of the first-stage effluent are used to estimate the first-stage feed quality (styrene content, bromine number and density). The first-stage reactor model integrates the estimated feed quality and measured reactor operating conditions, continuously inferring the firststage product quality: styrene, diolefins and bromine number. FIG. 10 illustrates the prediction of the styrene compared with online analyzer measurement. The second-stage reactor model integrates estimated feed quality and measured reactor operating conditions, continuously inferring the second-stage product quality. FIG. 11 illustrates the estimation of the BI in the effluent of the secondstage reactor. Using spot-detailed analyses and collection of operating conditions, APC users can generate the best tuning parameters to fit the current operation, thus allowing “realtime” control moves to improve performance.

1st stage optimizer Styrene inferred

Pygas inference

Reactor temperature optimal target Reactor quench ﬂow optimal target

Styrene Bromine index

MVAC multivariable controller FIG. 10. Styrene estimation in first-stage effluent by first-stage reactor model. Temperatures

Pressure

Reboiler

Flows

FIG. 12. APC architecture to optimize pygas hydrogenation.

BI inferred FC

H2 FC

Pygas feed

Quench ﬂow FC Diluent

Inlet temperature

Quench

Rx 1

ΔT B2 Styrene

TC = Manipulated variable Styrene = Controlled variable

FIG. 11. BI estimation in second-stage effluent by second-stage reactor model.

44 OCTOBER 2012 | HydrocarbonProcessing.com

Product to 2nd stage

FIG. 13. Simplified APC variables used for simulation example.

Process Control and Information Systems Recommended APC architecture. All APC components are embedded in an APC server connected to the DCS architecture, as depicted in FIG. 12. The control and optimization application consists of these modules: • MVAC module: State space multivariable predictive controller (MVPC) • Pygas inference • First-stage optimizer. The application provides one-minute cycles for MVAC (the MVPC) and 60-minute cycles for the optimizer. Controller execution time was determined by the process dynamics. FIG. 14. Eight-hour closed-loop APC simulation example.

APC PERFORMANCE Here is a simple example of APC potential, illustrated by the application to a real pygas unit optimization project. The control matrix components are presented in FIG. 13. The inputs or manipulated variables (MVs) are: • Feed flow to be maximized when available to reduce tank inventory • H2 flow used as long as available to prevent flaring • Reactor inlet temperature to control styrene content, but minimized when possible to lengthen catalyst cycles • Quench flow to control styrene content. The outputs or controlled variables (CVs) are: • Styrene in product, which should stay below the maximum limit • Reactor second-bed ΔT, which should stay below the maximum limit. When APC in turned ON, the styrene analyzer is at 1,300 ppm, below its 1,500-ppm maximum limit, and the reactor second-bed ΔT is at 65°C, below its 73°C maximum limit. As far as the operation is concerned, quench flow is too high, resulting in excessive cooling of the second bed. When the reactor inlet temperature is too high, a styrene giveaway occurs. More feed is available; the intermediate tank is not empty; and additional hydrogen is available, but is flared. APC actions on the process, as plotted in FIG. 14, can be summarized as making use of all available hydrogen to reduce styrene at the first-stage reactor outlet and thus reducing quench flow as far as the second-bed ΔT allows. These conditions will also reduce styrene content. Simultaneously, the reactor inlet temperature is reduced and feed flow is maximized within the constraints of the maximum styrene content limits. By better operation of the second bed, and using the 10% additional hydrogen available, this APC system was able to increase production by 10%, while decreasing reactor inlet temperature by 4°C. APC with inferential modeling has been successfully applied to pygas hydrogenation units. The overall benefits are: • On-specification product without giveaway • Ability to treat more feed : +10% • Reduction of first-stage reactor inlet temp.: –4°C • Catalyst run length maximization: +4 months/ current • Reduction of aromatics hydrogenation: –10%

• Reduction of H2 waste to flare: –10% • Energy savings: 5% The lengthening of the catalyst run length limits downtime for both the pygas unit and upstream units such as the steam cracker. An additional benefit observed by the operating staff was, with the ease of setting targets and the confidence that the APC system would meet these targets, they were free to concentrate on other plant activities. BIBLIOGRAPHY Bader, J.-M. and S. Guesneux, “Advanced process control optimization increases MTBE plant productivity,” Hydrocarbon Processing, October 2005. Bader, J.-M. and S. Guesneux, “Use real-time optimization for low-sulfur gasoline production,” Hydrocarbon Processing, February 2007. Tona P. and J.-M. Bader. “Efficient System Identification for Model Predictive Control with the ISIAC Software,” (ICINCO), Sétubal, Portugal, 2004. Grosdidier P. and J.-M. Bader, “Supervisory Control of an FCC unit through Sequential Manipulation,” Instrument Society of America, 1996. ACKNOWLEDGMENTS The authors express their gratitude to Joël Chebassier, Orionde, for his contribution to this article and also to the customers for making this article possible. This article is a revised and updated version from an earlier presentation at the American Fuels and Petrochemical Manufacturers (AFPM) Annual Meeting, March 11–13, 2012, San Diego, California. JEAN-MARC BADER is project manager at Axens’ Performance Programs Business Unit. His background is in energy engineering with over 23 years of experience in APC projects (design, development, implementation and maintenance) for refineries, petrochemicals and chemical plants (including ADU, FCC, CCR, alkylation, hydrotreating, ethylene, ammonia, blending), with several APC tools, on various DCS. He joined Axens in 2001, after several years with Elf and Total. His responsibilities include proposal and project management for APC projects. He graduated with honors from I.N.S.A. engineering school. GILDAS ROLLAND is a deputy product line manager— hydroprocessing and olefins for Axens. He started his career in 1998 at IFP Energies nouvelles as a process engineer in the R&D department. In 1999, he joined the process design department of the North American office in Princeton, New Jersey. In 2001, he moved back to Axens’ head-office where he served successively as start-up and tech service advisor, specialist in olefins technologies including R&D activities related to technology and catalyst improvement. Mr. Rolland was appointed to his current position in 2010. He is a graduate of the Ecole Centrale de Lille (E.C.Lille) and holds a master’s degree in refining and petrochemicals from the IFP School. Hydrocarbon Processing | OCTOBER 2012 45

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Special Report

Process Control and Information Systems M. J. KING, Whitehouse Consulting, Isle of Wight, UK

Why don’t we properly train control engineers? While there are managers in the process industry that see training control engineers as a “no-brainer,” these are very much in the minority. They may send staff on courses covering configuration of the distributed control system (DCS) and implementation of multivariable predictive control (MPC), but some managers seem to miss the point that engineers also need to develop expertise in basic control techniques. It appears to be a case of not knowing what they don’t know—i.e., there is a lack of appreciation of what a fully trained engineer can achieve. Without an injection of expertise, so-called “experienced” staff lack the knowledge to pass on to new recruits. Of the engineering disciplines relevant to the process industry, process control is probably the least well-taught at universities. Often handled by lecturers with backgrounds having little to do with chemical engineering, the courses are laden with complex mathematical techniques that have little relevance to the industry. While all graduates need additional training to advance their careers, this is particularly true for those destined to work in the field of process control. Process control engineers have an immediate impact on the process. Today’s systems permit the engineer to move from idea to commissioning with little involvement of other staff. Most other engineers develop recommendations that are reviewed with others, move on to designs that are also reviewed, and work with others during commissioning. Control engineers are more akin to process operators in the way they work. Operators are well-trained, so why aren’t control engineers? Questions to consider. The following 10 questions are designed to expose common gaps in a reader’s knowledge. If you are a control engineer, be honest in answering them: 1. Have all of the controllers been configured with the best choice of a proportional/integral/derivative (PID) algorithm? For example, am I aware that most systems support the option to have proportional action based on the process variable (PV), rather than on error? Do I believe that this algorithm is inferior because it gives a slow response to setpoint (SP) changes, or do I know that, for many controllers, applying this option with the correct choice of tuning can reduce, by a factor of three, the time that it takes the process to recover from a disturbance? (See FIG. 1.) 2. Am I using trial-and-error as the main tuning method? Am I aware that this increases, by a factor of around 50, the time taken to properly tune a controller? Do I know that, because of the time required, the controller is unlikely to ever be properly tuned? Am I aware that there are over 200 tuning methods published for PID control, and that most—if not all—of them

have some major deficiency? Does my chosen method properly compromise between a fast return to SP and the movement of the manipulated variable (MV)? (See FIG. 2.) Is this method designed to be used with the chosen version of the PID algorithm? 3. Do I know that applying derivative action can greatly improve controller performance if the process deadtime is large compared to the lagtime? (See FIG. 3.) Am I reluctant to use it because it makes tuning more complicated? Do I abandon its use if the measurement is noisy, or do I know how to solve this problem? Do I know how to resolve the spiking problem that derivative action causes with regard to discontinuous signals? 4. Is maximum use made of the surge capacity in the plant? (See FIG. 4.) Are vessel levels maintained close to SP, or are they allowed to approach alarm limits to minimize downstream flow disturbances? Are level gauges ranged to maximize vessel working volume? Do I know that nonlinear algorithms such as “error squared” and “gap control” can be used to more fully exploit surge capacity? SP PV (proportional-on-error) PV (proportional-on-PV)

FIG. 1. Response to a load change.

PV (limiting MV overshoot) PV (ignoring MV) SP MV (acceptable overshoot) MV (unacceptable overshoot)

FIG. 2. Taking account of MV overshoot. Hydrocarbon Processing | OCTOBER 2012 47

Process Control and Information Systems 8. Do I apply density compensation to fuel gas flow controllers to display flowrates in standard volumetric units (e.g., Nm3/ hr or standard cubic feet per minute)? Do I know that this worsens the disturbance caused by changes in gas heating value? 9. Are my inferential property calculations automatically updated using laboratory data? Am I aware that, in most cases, this can cause the inferential to become less accurate? 10. Have I been persuaded to locate my compressor controls in specialist hardware rather than in the DCS? Do I know that, if I apply the correct tuning method, this may not be necessary? How did you do in the test? If it has exposed even one area where your knowledge is incomplete, then chances are that there is an opportunity to improve process performance that will capture benefits far excost ceeding the cost of effective training.

5. Are filters being used mainly to reduce the visual impact of noise on trended variables? Filters can significantly reduce the controllability of the process and may not be necessary in all cases. Do I know that I should instead check what impact the noise has on the final control element (usually a control valve)? Do I know of other readily available filtering techniques that cause less distortion to the base signal? Am I aware of the importance of eliminating noise at the source, particularly with level measurements, and how this can be achieved?

Winning even one more contract by demonstrating a higher level of expertise more than justifies the of developing that expertise.

6. Am I aware of other algorithms that can outperform even an optimally tuned PID algorithm? Do I know that these can be easily implemented in most DCSs? 7. Do I know that most MPC packages provide bias rather than ratio feedforward? In many cases, performance can be substantially improved by implementing ratio feedforward at the DCS level. Do I know how to properly tune the dynamic compensation in such controllers? Do I know of the benefit that ratio feedforward gives in automatically maintaining optimum PID tuning in all of the unit’s controllers as the feed rate is changed?

SP PV (PID) PV (PI)

FIG. 3. Use of derivative action.

Averaging control Tight control Vessel level

Downstream ﬂow

FIG. 4. Use of surge capacity.

48 OCTOBER 2012 | HydrocarbonProcessing.com

Training costs. What does it cost to train a control

engineer, and what are the economic benefits? In addition to the time spent on learning how to configure the DCS and how to apply the chosen MPC, a control engineer will need around three weeks of further training. This training should cover basic control techniques, “conventional” advanced control, process-specific techniques, inferentials, etc. Such courses can cost $1,000 per day. Factoring in travel and living expenses, the total price of training could be $20,000. A manager might view this as costly, but it is insignificant compared to the benefits to be achieved through additional training. For example, a control engineer typically will be responsible for control applications that are capable of capturing in excess of $500,000 per year. Commissioning a project of this value just two weeks sooner would be enough to justify the training. If maintaining existing applications (for example, over a two-year period), then a 2% increase in their utilization would generate the same savings. Also, if the company relies on external specialists during implementation, then reducing the involvement of a top-grade consultant by two weeks would yield similar savings. While such benefits apply to operating companies, similar benefits can be achieved by those companies offering advanced process control (APC) implementation and process engineering services. With only minor differences between competing technologies, the main criterion in selecting an APC implementation company is the expertise of the engineers it offers. Winning even one more contract by demonstrating a higher level of expertise more than justifies the cost of developing that expertise. Similarly, plant owners are increasingly expecting engineering contractors to be more aware of the importance of good basic control design. Too many processes with inherent control problems exist, along with missed opportunities that could have been avoided at negligible cost, if considered at the process design phase. Which course should an engineer choose? More than any other engineering subject, process control training requires practical, “hands-on” exercises. Most engineering disciplines work with steady state. It is relatively easy to demonstrate steady-state behavior in a computer slide presentation. However, it is not so easy to show parameters changing over time.

Process Control and Information Systems Student-friendly, dynamic simulations take far more time to build; it can take 50 hours or more to develop the material covered in one hour on the course. The ratio for the preparation of more conventional teaching material is likely less than 10:1. More effective courses are necessarily more costly. This is particularly true if they are presented by the more experienced— and, therefore, usually more highly paid—engineer. The value of a course should be assessed on what impact the participant can have on process profitability upon returning to work. He or she should return with several ideas that can be put into practice immediately. Presenting the course on a manufacturing site provides the opportunity for practical exercises to be carried out on real controllers. The resulting improvements have a noticeable impact on process performance, and they greatly increase the confidence of the engineer to implement other ideas. Who should present the course? It might be easier to answer this question by identifying potentially poor choices. The DCS vendor is best placed to instruct staff in the use of the system. However, vendors are generally more effective at explaining the “how” than the “why.” For example, they can describe the multiple versions of the PID algorithm available in their systems, but they are generally less adept at explaining when each algorithm should be used. Similarly, the MPC suppliers will be able to describe how to effectively design, implement and monitor their technology, but they will not go into detail about the basic controls that should be in place before step-testing is undertaken. While MPC suppliers are concerned that such controllers operate well, they generally place less demanding criteria on their performance. With a few notable exceptions, most academic institutions treat process control as a highly theoretical subject. Their courses tend to be cheaper because the tutor’s time and the facilities have already been paid for; however, their usefulness is often questionable. Should the course be held in-company? There is the

temptation, particularly if only one or two engineers need training, to send them on an open-access course. It costs the supplier more to run these types of courses than it does to run in-company courses since open-access courses must be marketed to a wide client base, there is a greater administrative load, and the course facilities must be rented. For the customer, an open-access course may be the less costly option, even with the inclusion of travel and living expenses. Also, engineers may have the opportunity to develop valuable contacts in other organizations. However, the following points should be considered: • An in-company course opens up the opportunity for others to attend; the most successful APC projects are those in which the entire staff is involved. • Plant supervisors, process engineers and production planners normally do not attend open-access process control courses; however, they will usually sit in on at least part of an in-company course. An in-company course provides a valuable opportunity for these personnel to develop an awareness of technology and the role they can play in its successful implementation.

• An in-company course can be customized to closely match the company’s needs. • Some material included in an open-access course may not be relevant; it may assume less previous knowledge, and its timing may be inconvenient. When should training take place? Training budgets, like

many expenses that are perceived as optional, are often the first to be cut when the economic climate is poor. However, this is precisely the time when control engineering expertise should be developed. The likelihood is that no major APC projects will be approved, and so releasing engineers for training does not disrupt their schedules. Furthermore, engineers will have time to identify and exploit the many zero-cost improvements revealed by the training. Also, when major investments are again considered, the basic process control layer will already be ready to receive APC— therefore, substantially shortening its commissioning. MYKE KING is the author of Process Control: A Practical Approach, as well as the director of Whitehouse Consulting. Previously, he was a founding member of KBC Process Automation. Prior to that, Mr. King was employed by Exxon. He is responsible for consultant services, assisting clients with improvements to basic controls, and with the development and execution of advanced control projects. Mr. King has 35 years of experience in such projects, having worked with many of the world’s leading oil and petrochemical companies. He holds an MS degree in chemical engineering from Cambridge University, and he is a Fellow of the Institution of Chemical Engineers (IChemE).

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Special Report

Process Control and Information Systems A. J. SZLADOW, REDUCT & Lobbe Technologies Inc., Richmond, British Columbia, Canada

Consider automated fault detection systems to improve facility reliability Automated fault detection and diagnosis systems (AFDDSs) are well established in many consumer and industrial sectors.1 The conventional limit-value based (high/low alarms) fault detection and diagnosis systems have the advantage of simplicity and reliability. Yet, they also have a major weakness. These systems can only react to the deterioration of system conditions, and they do not provide sufficient time and information to detect and diagnose anomalous conditions proactively. TABLE 1 summarizes the relative advantages of AFDDS vs. standard FDDS control. This article addresses how to implement an AFDDS in a refinery, and discusses the advantages and key issues with AFDDS.

AFDDS APPLICATIONS There are numerous publications on fault detection and diagnosis in electrical systems, including application of statistical and soft computing methods. However, very little of the knowledge and experience gained from AFDDS application from other industrial sectors has been applied to the refining industry. A literature review regarding methods applied in AFDDS in heavy industry lists 367 references.1 None of the 367 listed references refers to applications found in petroleum refining. In 1995 and 1997, similar literature reviews identified over 250 applications of intelligent systems including, AFDDS in heavy-industry operations.2 Again, the majority of the applications addressed process control and optimization, scheduling, and design for productivity and product quality. Less than 10% of the applications described fault detection and diagnosis systems, and most cases were not automated; they primarily provided decision support to process operators and engineers. The number of AFDDSs applied appears to be related to the automation level of the site plant and the risks associated with unsafe operating conditions. The aerospace sector has a very high level of AFDDS applications due to the risk associated with this sector and corresponding levels of regulation. Given the listed considerations and that very few AFDDSs have been implemented in oil refineries, we will discuss AFDDS application in petroleum refining and also review systems implemented by industry to manage process and equipment failures.

of $7.5 million with the application of predictive diagnosis in critical refinery process units.3 An EPRI report examined present adoption of control technologies in California refineries and the move to use of distributed control systems (DCSs), multivariable, neural networks and future self-learning tools as shown in TABLE 3.4 Such progression of technological changes has a large potential benefit. But, this progression will require investment. Following the methodology outlined in the EPRI report, Table 4 shows the avoided maintenance cost for refineries “before and after” implementation of advanced control technologies. We assign any reduction in maintenance costs to AFDDS. TABLE 4 depicts three levels of technology implementation: 1. Present level with advanced technologies 2. Application of marketable (present available cost-effective) technologies 3. Application of “technical potential” (future cost-effective) advanced technologies. A large difference (double) between marketable and technical potential technologies indicates possible gains in technology capacity through research. Cost reduction. The maintenance cost was estimated at 7% of

the total operating cost (Salomon 2006), which yields a benefit of over $1.3 million/yr from increased penetration of available technologies and a further $1.7 million/yr from implementing future (hybrid and self-learning) technologies (see TABLE 4). The annual total reduced maintenance cost is estimated at over $3 million for an average refinery. TABLE 1. Standard controls vs. AFDDS Issue

AFDDS

Alarm detection

Reactive

Proactive/Prognosis

Alarm management

Non-deductive

Deductive notiﬁcation

Personnel guidance

Little

Signiﬁcant

Inputs

Largely from sensors

Sensors and expert knowledge

Corrective action

Operator initiated Automatic

Automatic

Used by

Operating personnel

Operating, maintenance, engineering and safety personnel

Examples. There are very few published references to the ben-

efits of AFDDS within the petroleum sector, as summarized in TABLE 2. A report by Berra indicates that one client gained savings from the reduction of unplanned shutdowns on the order

Standard control

Hydrocarbon Processing | OCTOBER 2012 51

Process Control and Information Systems TABLE 4. Avoided maintenance costs of Canadian reﬁneries depending on the level of automation

TABLE 2. Summary of methods used and process units studied Reference Wang11

MPC

PCA

Vedan12

Huang13

FCC

Yang

Yamamoto15

Process Distillation column

Pranatasta

Coker v

FCC

FCC v

Pranatyasto16

CC v

FCC v

Gofuku Du18

Reﬁnery

v 19

Wilson

FCC

ES—Expert systems MPC—Model predictive control NN—Neural networks

Utility PCA—Partial component analysis FL—Fuzzy logic SI—Semantic interface

TABLE 3. Present state of adoption of control technology in California reﬁneries Sub-section of the reﬁnery, %

Whole reﬁnery, %

Move to DCS

Move to multivariable

Move to neural network

Future self-learning control

Present control technology

Current level with the use of advanced control technologies

Avoided cost, $ million 0

Application of “marketable” (cost-effective today) advanced control (AFDDS) technologies

> 1.3

Application of “technical potential” (cost-effective in the future) advanced control (AFDDS) technologies

> 1.7

FCC

Patan

Level of automation

Source: EPRI, 20044

The lost production due to unscheduled shutdowns is typically reported to be between 10% to 20% of operating costs.5 Assuming a potential 50% reduction in unscheduled shutdowns, etc., from installing an AFDDS, the estimated annual benefits would exceed $3 million for an average refinery. However, to achieve this level of benefits, an AFDDS would have to be applied on half of the following unit operations and processes: 1. Unit operations: Distillation, absorption columns, furnaces, heat exchangers and compressors. 2. Unit processes: Fluid catalytic cracking (FCC), catalytic reformer, hydrocracking, delayed coking, hydrotreating and alkylation. About six AFDDS models would be required to gain the benefits listed.

AFDDS PERFORMANCE Heavy industry has used advanced process control (APC) systems for optimization projects and has given much less attention to AFDDSs. However, depending on the level of automation, benefits from managing abnormal process and equipment conditions can increase reliability; the benefits often exceed the gains found through process optimization. For example, for a large continuous operation, such as a refinery, process optimization can typically yield a 3% improvement in productivity.6 In contrast, a well-implemented AFDDS may yield up to a 5% improvement in profitability. This is because management for reliability improvement goes beyond fault avoidance by providing the ability to: 52 OCTOBER 2012 | HydrocarbonProcessing.com

1. Handle large disturbances and control variables at their optimal values 2. Ensure and upgrade dynamic process models, including factors omitted in initial implementation 3. Explain the behavior of controllers and, when needed, correct controllers to meet planned targets 4. Provide advice on alarm management, including early detection of problems before more serious problems develop. As summarized in TABLE 5, it is possible to classify fault detection and diagnosis methods into quantitative using models based on first principles, qualitative using models describing lumped system responses, or process history methods matching fault patterns derived from historical data.7–9 The methods are similar and, yet, different from each other. They can identify the relative strengths and weaknesses from methods when building diagnostic methods for fault detection and diagnosis (anomaly detection, disturbance detection and controller diagnostics) and supervisory control (controller tuning, control reconfiguration and online optimization). AFDDS not only address typical maintenance functions such as better root-cause analysis or optimized inspection frequencies, but, in 7 out of 10 cases, they also address processing issues. Safety and reliability. Improving operational safety and meeting regulatory requirements are critical to industry operations and businesses. For example, AFDDSs have been applied for safety and regulator reasons in the automotive and aerospace sectors. This article does not discuss using AFDDSs to enhance safety and meet regulatory requirements in refineries. It is assumed that, where needed, the refining industry would implement such systems as required. Higher reliability due to AFDDS results in a more energy efficient and profitable facility. However, AFDDS-driven energy savings are often indirect through less production waste, reduced plant outages, less plant startups and/or shutdowns, and more optimal equipment/process performance through better control systems management, etc. All lead to reduced net energy consumption per product unit made, or higher overall plant energy efficiency.

RELIABILITY FOCUS A simple focus on benefits/cost analysis does not reflect the true opportunities created by AFDDS technology. For such an analysis, it may not include technologies that can be adopted easily and will later lead to significant learning and a significant cost reduction. Therefore, broader adoption criteria, such as those listed in TABLE 6 can provide better guidance as to the best AFDDS projects and development directions to support.

Process Control and Information Systems TABLE 5. Summary of fault detection and diagnosis methods Quantitative methods

Qualitative methods

History methods

Given required measurements can distinguish known from unknown faults

Can provide explanation for fault propagation

Fault rules can be used where fundamental principles are lacking

Can detect faults for systems with process and measurement noise

Can generate and recognize full set of faults

Have been demonstrated to perform well in terms of robustness to noise and resolutions of parameters

Effectiveness is determined by sensor data and system knowledge

May have poor resolution due to ambiguity of qualitative reasoning

Easy (time and cost) to implement

Approximation of disturbances can create modeling errors

Resolution problems can be addressed with quantitative information

Poor fault generalization from historical data only

Complex systems modeling may generate spurious solutions because of computational complexity

May have difficulty with multiple faults depending on algorithm used Limited by a ﬁnite set of data

TABLE 6. Summary of potential AFDDS beneﬁts and costs, millions of dollars

TABLE 7. Examples of industry-wide AFDDS adoption criteria Overall

Cost element

Detail

Must represent a forward step

Potential for learning by doing and/or research

Average beneﬁt

Average cost1

Reduction in maintenance cost

1.4

1.1–2.1

Cross-cutting potential

Reduction in cost of outages

3.1 1.1–2.1

Customization requirements

Plants

Total reduction and cost Beneﬁt/cost ratio 1 2 3

Comments

Petroleum

4.5

2.1–4 (1.5–2.8)3

Have clear adoption AFDDS barriers issue Host plant expertise

Based on six AFDDSs models per plant Inline with $4 to $6 million reported by Stout21 and Kant20 Based on reduction in cost of outages only

However, in each technology stage, there are niches or specialized markets that often experience adopting new technology much sooner. Some general criteria for applications of AFDDSs in oil refineries sector should be sought, explored and emphasized to commercialize AFDDSs at a faster rate in areas such as FCC. Implementing an AFDDS requires a high investment in knowledge of refining operations, plant controls and automation, information technologies and software, advanced technologies for data analytics and visualization, plant-wide information systems, etc. The market size of the petroleum refining industry, therefore, is critical for the private sector to justify investment of the large amount of resources and manpower required. A more detailed discussion of the barriers to the introduction of advanced control technologies in the refinery sector and other heavy industry sectors can be found in the literature.10 The progress in AFDDS is not likely to come from a large breakthrough in science and technology, but from incremental improvement in the cost of AFDDS and the gradual acceptance by industry. TABLE 6 shows the estimated potential benefits and costs for an AFDDS, and a few observations are evident: 1. Because of the large difference between an average AFDDS cost (about $200,000) and the cost of improving plant digital infrastructure ($500,000 to $1.5 million), it is the plant’s existing infrastructure (or required improvement in infrastructure) that drives the benefit/cost ratio for implementating an AFDDS at any plant/refinery. 2. Assume that in all AFDDS cases, some infrastructure will have to be updated. In spite of that, in all cases, benefit/cost ra-

Lack of support/ sponsorship

Doesn’t have to be new AFDDS technology

Clear and doable Technology is easy— People are hard

Securing technical expertise Business ownership Innovative ﬁnancing options Supports areas of major interest

Relevance or impact on oil reﬁneries sector Proponent expertise A niche application

Identiﬁcation of applications relevant to the sector’s strategic investments will accelerate AFDDS adoption and increase capacity

tios of > 0.4 (with approximately 1 most likely) and paybacks of less than two years (about one year most likely) were projected. In the final analysis of AFDDS adoption, one has to ask: What if the oil refining industry or a large industrial segment does not adopt AFDDS technologies? It is difficult to predict the future of a specific industry, but strong conclusions can be made based on what is known about the role of technology in industry growth: 1. Technical progress is the most important factor in economic growth, and, typically, it accounts for more than half of growth in developed countries. 2. Industries that use advanced technology are more productive and profitable and have higher wages. 3. Industries that use advanced technologies have higher job growth. 4. New technologies revitalize old industries, e.g., steel, automobile, textile, etc. Hydrocarbon Processing | OCTOBER 2012 53

Process Control and Information Systems Also, industry-wide AFDDS adoption criteria can be formulated, as shown in Table 4. Failure to implement AFDDS or slow progress to adopt is likely to result in a loss of opportunities as measured by productivity within the petroleum refining industry. LITERATURE CITED Chiang, L. H., E. L. Russel, and R. D. Braatz, Fault detection and diagnosis in industrial systems, Springer, 2001. 2 REDUCT and Lobbe, Technologies, “Application of Intelligent Systems to increase productivity, quality and energy efficiency in heavy industry,” and “Advances in the application of Intelligent Systems in heavy industry,” CANMET Technology 1995 and 1997. 3 Berra, J., “The digital refinery: A look at the future of automation,” NPRA Computer Conference, 2002. 4 EPRI Report 1007415, Using advanced control and power technologies to improve the reliability and energy efficiency of petroleum refining and petrochemical manufacturing in California, 2004. 5 White, D., “The 21st century refinery: Impact of modeling and predictive analytics,” NPRA Technical Forum on Plant Automation, 2007. 6 Gosh, A. and D. Wall, “Abnormal conditions management–The missing link between sustained performance and costly disruptions,” ARC Advisory Group, March 2001. 7 Venkatasubramanian, V., R. Rengaswamy, S. N. Kavuri and K. Yin, “A review of process fault detection and diagnosis, Part III: Process history based method,” Computers & Chemical Engineering, 2003. 8 Venkatasubramanian, V., R. Rengaswamy and S. N. Kavuri, “A review of process fault detection and diagnosis, Part I: Quantitative model-based methods,” Computers & Chemical Engineering, 2003. 9 Venkatasubramanian, V., R. Rengaswamy and S. N. Kavuri, “A review of process fault detection and diagnosis, Part II: Quantitative models and search strategies,” Computers & Chemical Engineering, 2003. 10 Szladow, A., “Developing intelligent systems for heavy industry: The adoption of 1

intelligent technologies,” PCAI, Vol. 17.6, 2005. Wang, X. Z., et al., “Learning dynamic fault models based on a fuzzy set covering method,” Computers & Chemical Engineering, Vol. 21, No. 6, 1997. 12 Vedam, H. and V. Venkatasubramanian, “PCA-SDG based process monitoring and fault diagnosis,” Control Engineering Practice, Vol. 7, No. 7, 1999. 13 Huang, B., et al., “Fault diagnosis of an industrial CGO coker model predictive control system,” IEEE Canadian Conference, 1999. 14 Yang, S. H., B. H. Chen and X. Z. Wang, Engineering applications of Artificial Intelligence, Vol. 13, No. 3, 2000. 15 Yamamoto, J., et al., “Application of a cooperative control system to residue fluid catalytic cracking plant using a knowledge based system and model predictive multivariable control,” IECON 2000. 16 Pranatyastos, T. and S. J. Qin, “Sensor validation and process fault diagnosis for FCC units under MPC feedback,” Control Engineering Practice, Vol. 9, No. 8, 2001. 17 Gofuku, A., and Y. Tanaka, “Display of diagnostic information from multiple viewpoints in anomalous situation of complex plants, systems, man and cybernetics,” IEEE International Conference, 1999. 18 Du, D., et al., “Expert System for diagnosis and performance of centrifugal pumps,” 1996. 19 Wilson, D., A. Jumenez and J. Souza, “An on-line advisory system for optimizing refinery utilities systems,” NPRA Technical Forum on Plant Automation, 2006. 20 Kant, R., and K. Pihlaja, “Abnormal situation prevention (ASP) in complex system,” NPRA Plant Automation Conference, 2006. 21 Stout, J., “Reliability and operations management applications in olefins plants,” AIChE Spring National Meeting, Houston, April 2001. 11

ADAM J. SZLADOW is president of REDUCT & Lobbe Technologies. He has over 30 years of experience in the development and application of advanced technologies in energy and heavy industry. He held management and research positions in utility industry, energy development companies and government research laboratories. Dr. Szladow was chairman of the Business Committee of the National Advisory Council to CANMET, Natural Resources Canada; and a member of the Minister’s National Advisory Committee, Natural Resources Canada. He holds a PhD in materials sciences and chemical engineering from Pennsylvania State University, and has authored over 70 scientific publications including patents.

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Process Control and Information Systems

Special Report

V. YADAV, P. DUBE, H. SHAH and S. DEBNATH, Indian Oil Corp., Ltd., Mathura, Uttar Pradesh, India

Optimize desulfurization of gasoline via advanced process control techniques At Indian Oil Corp.’s (IOC’s) Mathura refinery, a selective desulfurization unit was commissioned to reduce the sulfur content of fluidized catalytic cracked (FCC) gasoline—a blending component for finished motor spirit (MS). The objective of this new unit was lowering the sulfur content of FCC gasoline from 500 ppmw to 100 ppmw, thus meeting Euro IV product specifications for the refinery-gasoline blending pool. However, along with desulfurization, some undesirable olefin saturation reactions occurred, resulting in octane losses for the product gasoline. As per design, the octane loss in the desulfurization reactors is 1.3 units. With Euro IV specifications in place, the octane loss negatively impacted the refinery’s economics. This refiner applied an advanced process control (APC) solution to minimize octane loss. The objective of the desulfurization unit’s APC program is to maximize sulfur content in the gasoline while still complying with Euro IV specifications and other process operating constraints. The control philosophy depended on sulfur estimations of the stabilizer-bottom product. An inferential property was developed for online estimation of the sulfur content, and it was used as a controlled variable in the multivariable predictive controller (MVPC). This case history describes the development of the inferential models used in open-loop and closed-loop applications, laboratory and analyzer update mechanisms, and APC model generation. With APC, it was possible to increase the sulfur content in product gasoline by 10 ppm–12 ppm, along with an average octane gain of 0.11 units; all improved the Flow of SHU refinery’s bottom line.

reactor effluent is separated into light-cut naphtha (LCN), heart-cut naphtha and heavy-cut naphtha (HCN) in the FCCGS unit. In the third step, the heavy fraction from the splitter bottom, containing high-sulfur content material, is processed in the HDS unit. This processing step converts heavy sulfur compounds into hydrogen sulfide (H2S). In addition, significant saturation of olefins occurs along with the HDS reactions. Saturating olefins reduces the final research octane number (RON) and is an undesirable condition.

ADVANCED APC OBJECTIVES AND DESIGN In the Mathura refinery application, the control objectives are achieved by utilizing MVPC in conjunction with supporting predictions provided by an inferential property prediction package (IPPP). Supporting calculations are required to supplement existing process measurements. MVPC applications incorporate process models that permit forward-feed disturbance rejection and intermediate variables feedback, as well as constraint control. In configuring the controller, there is one main controller. The objectives for the main controller are: • Maximizing stabilizer-bottom product sulfur level within permissible limits, so that the upper limit of the total rundown sulfur for the desulfurization unit is maintained per Euro IV gasoline blending. Minimizing RON loss is also achieved. • Minimizing steam consumption by the stabilizer section • Maintaining safe unit operations.

gasoline recycle

FCC GASOLINE DESULFURIZATION PROCESS IOC’s Mathura refinery implemented a new gasoline desulfurization process. It is a two-step selective hydrotreating method. This processing unit consists of three major operations: • Selective hydrogenation unit (SHU) • FCC-gasoline splitter (FCCGS) unit • Hydrodesulfurization (HDS) unit. In the first step, FCC gasoline is treated in the SHU, which selectively converts di-olefins into olefins and light mercaptans into heavier sulfur-containing compounds. In the second step, the SHU

306FIC0101 FCCU debutanizer ﬂow

306-E-02 306TIC0270

306-E-01B

306FI0105

307-R-01

306LIC103

306-V-01

306FIC0104

306-E-01A

Steam

306TI0263 306-V-04

306FIC0203 To FCCGSU gas feed

Manipulated variables Controlled variables Disturbance variables

306-P-01A/B 306FIC0202

306-V-05

Controller: MainCON Sub-Controller: SHUCON

To FCCGSU liquid feed

FIG. 1. Block diagram of the sub-controller for the selective hydrogenation unit. Hydrocarbon Processing | OCTOBER 2012 55

Process Control and Information Systems To achieve these objectives, a main controller (MAINCON) and two sub-controllers are used: • Selective hydrogenation unit—SHUCON • Hydrodesulfurization unit—HDSCON. Note: The stabilizer section of the HDS unit is considered part of HDSCON.

ering the debutanizer flow (hot feed) as a DV. The SHU feed/ effluent exchanger bypass flow, along with steam to the SHU preheater, is used to control the SHU reactor-inlet temperature. FIG. 1 shows the same sub-controller (SHUCON) for the SHU. HDS unit sub-controller. Before the APC implementation,

the HDS unit was operated by controlling the severity conditions of the reactors. The unit operator conThe objective of the desulfurization unit’s trolled HDS reaction (first-bed inlet temperature and second-bed inlet temperature) based on daily sulfur APC program is to maximize sulfur content levels in the stabilizer bottom product and rundown in the gasoline while still complying with Euro IV. product. Sulfur levels were determined by analyzers and lab testing. The fuel gas was cascaded with firstbed inlet temperature, and the quench flow was cascaded with second-bed inlet temperature. To maintain stable Sub-controller objectives. Before the APC installation, reflux flow to the stabilizer, unit operators adjusted the stabithe SHU was operated to maintain stable flow to the reactor. lizer reboiler temperature and reflux pressure by continuous Flow from the FCCU debutanizer (hot feed—70% of total) monitoring of the light-end flow to the column. was routed to a feed-surge drum. A recycle stream (HDS stabilizer bottom stream) from a nitrogen-blanketed storage (cold feed—30% of total) was also sent to the feed-surge drum. Unit Post-APC operations. The HDS reactor is set by the APC operators manually controlled the level of the SHU feed-surge based on sulfur levels of the stabilizer bottoms. The IPPP estidrum by adjusting the recycle stream. mation is done on a 15-second basis. Also, the APC will maxiTo maintain the SHU reactor inlet temperature, feed from mize the sulfur level within given operator limits, thereby by the surge drum is heated by the SHU feed-effluent exchanger adjusting the reactor severity. The stabilizer-bottom reboiler on the tube side by exchanging heat from the SHU effluent. The temperature is controlled by APC and facilities minimizing the resulting mixture is heated in the SHU preheater using steam. steam consumption by the reboiler. However, the reflux flow to After the APC installation, the control objective was to the stabilizer is also controlled by APC, along with stabilizerkeep steady flow to the SHU feed and maintain the surge-drum bottom re-boiler temperature. The process equipment manlevel by adjusting the FCCU debutanizer flow as a disturbance aged via APC includes: variable (DV) and adjusting the recycle stream. The control • HDS reactor (307-R-01) objective is to maintain a stable SHU RIT, by manipulating the • HDS heater (307-F-01) effluent exchanger bypass flow and steam flow to the SHU pre• HDS feed-effluent exchanger (307-E-01 A/B/C/D) heater under allowable limits. The process equipment to be • Stabilizer section (307-C-02). managed via the APC included: TABLE 2 summarizes the sub-controller design for the HDS • SHU feed-surge drum (306-V-01) unit. FIG. 2 shows the same sub-controller (HDSCON) for the • SHU feed-effluent exchanger (306-E-01A/B) HDS unit. • SHU preheater (306-E-02). TABLE 1 summarizes the sub-controller design for the selecModels. As shown in FIG. 3, the simple first-order process models tive hydrogenation unit. The SHUCON sub-controller was dewere not providing tight control on the HDS reactor-inlet temsigned to manage steady flow to the SHU reactor while considperatures. In response, a ramp transfer function block was added into the model, along with the first-order transfer function block. The exothermic Steam 307TI0607.PV 307-E-06 reaction in the reactor behaves in a “ramp” 307TI0642.PV manner (unbounded runaway even in the 307-V-04 case of a bounded input disturbance). 307FIC0605.SP 306FIC0502.PV Due to “ramp” behavior of the process, fast 307-E-01 action is required in manipulated variables 307FIC1003.PV 307-R-01 (MVs), such as fuel-gas flow and quench 307-E-01 307TIC0635.PV flow, to quickly control the exotherm (by 307-C-02 307-F-01 controlling the first-bed and second-bed 307-E-01 inlet temperatures) before they rise too 307-E-04 Manipulated variables 307-E-05 307TI0630.PV Controlled variables high. The inherent instability of the reacDisturbance variables tor was countered via a ramp block, plus 307TI1014.PV Controller: MainCON Sub-Controller: HDSCON the normal first-order block, to relate the STABBTM_SULFUR PV(RQE) MVs and DVs with the inlet temperatures. 307FIC0684.SP 307-V-06 307PIC1003.SP For a step change in DVs, this combina20TI0804.PV 307FI0606.PV To rundown 307AI1001.PV tion predicts an unbounded response in the inlet temperatures—thus, moving the FIG. 2. Block diagram of the sub-controller for the hydrodesulfurization unit. MVs quickly to reject the disturbance. 56 OCTOBER 2012 | HydrocarbonProcessing.com

Process Control and Information Systems SUPPORTING CALCULATIONS IPPP DEVELOPMENT To calculate the sulfur content of the FCC feed inlet, several predicted values were considered. By using the flowrate and sulfur quantity of all streams listed in TABLE 3, the total sulfur value can be calculated at the FCCU feed inlet. The calculation used to estimate the sulfur content is: =(79FC803.PV S1 density) + (79FC802.PV S2 density) + (79FC801.PV S3 density) + (7FC6701.PV S4 density) + (12FIC100.PV S6 (if crude_select.op=1) density) or (MRA.12FIC100.PV S7 (if crude_select.op = 2) density) or (MRA.12FIC100.PV S8 (if crude_select.op = 3) density) + ((2FC0708.PV S5)/1000) / (79FC803.PV + 79FC802.PV + 79FC801.PV + 7FC6701.PV + 12FIC100.PV + 2FC0708.PV) where S1–S8 are sulfur values that are entered by the operator.

FCCDSU feed sulfur. This model used several inputs:

Tag name FCCUFD_SULFUR.PV 19TRC153.PV

Tag description Sulfur at FCCU (calculation) FCCU main fractionator top temperature 20TI99.PV FCCU debutanizer bottom temperature. To estimate the sulfur content of DSU feed, the following linear equation is used: P = Ax1 + Bx2 + Cx3 + Bias where: P = DSU_SULFUR.PV (FCCDSU feed sulfur in hot feed) A = Coefficient 0.041417 x1 = FCCUFD_SULFUR.PV B = Coefficient 1.6497 x2 = 19TRC153.PV C = Coefficient 5.736500 x3 = 20TI99.PV Bias = –1067.4

Sulfur content of FCC gasoline splitter. Feed to FCCGSU

is compensated by two streams—hot feed from the debutanizer (306FI0105.PV) and cold feed from recycle (306FIC0101. PV). Calculations to estimate sulfur at FCCGSU feed are: = ((DSU_SULFUR.PV 306FI0105) + (STABBTM_ SULFUR 306FIC0101.PV-5.5*)) / {(306FI0105) + (306FIC0101-5.5*)} where DSU_SULFUR.PV and STABBTM_SULFUR are the IPPP sulfur estimations. * 5.5 is the flow correction since the control valve has a zero error. IPPP applications. Several IPPP models were developed for

the FCC gasoline desulfurization unit and include: • FCCDSU hot feed sulfur estimation • HDS feed sulfur estimation • Stabilizer bottom sulfur estimation.

FIG. 3. First-order process model response to reactor inlet temperature control.

TABLE 1. APC variables for the sub-controller for the selective hydrogenation unit—SHUCON Description

Interface point

Manipulated variables: Flow of SHU gasoline recycle

306FIC0101.SP

SHU feed/effluent excahnger bypass

306FIC0202.SP

Steam ﬂow to SHU pre heater

306FIC0203.SP

Disturbance variables: FCCU debutanizer ﬂow

306FI0105.PV

Flow to SHU from surge drum

306FIC0104.PV

Controlled variables: Feed surge drum (306-V-01) level

306LIC0103.PV

SHU reactor (306-R-01 B) inlet temperature

306TIC0270.PV

SHU pre-heater inlet temperature

306TI0263.PV

FIG. 4. Quality and process improvement achieved through APC IPPP. Hydrocarbon Processing | OCTOBER 2012 57

Process Control and Information Systems TABLE 2. APC variables for the sub-controller for the HDS unit— HDSCON Description

Interface point

Manipulated variables: Fuel gas ﬂow

307FIC0684.SP

HDS reactor 2nd bed quench

307FIC0605.SP

Stabilizer bottom steam pressure

307PIC1003.SP

Disturbance variables: HDS feed from GSU

307FI0606.PV

Stabilizer light end feed from GSU

306FIC0502.PV

HDS reactor 2nd bed bottom temperatue

307TI0630.PV

HDS feed temperature at GSU

20TI0804.PV

HDS feed sulfur

HDSFD_SULFUR.PV

Controlled variables: HDS reactor 1st bed inlet (307R01) temp

307TI0642.PV

HDS reactor 2nd bed inlet (307R01) temp

307TIC0635.PV

Feed effluent exchanger inlet temp

307TI0607.PV

Stabilizer (307-C-02) bottom temp

307TI1014.PV

Reﬂux ﬂow to the stabilizer

307FIC1003.PV

Online stabilizer bottom sulfur

307AI1001.PV

Stabilizer bottom sulfur (inferred)

STABBTM_SULFUR.PV

HDS feed sulfur. This model used several inputs:

Process inputs used Tag name Tag description GSUFD_SULFUR.PV Feed to FCCGSU (calculation) 20PI0802.PV FCCGSU top pressure 20FC0306.PV FCCGSU light cut draw flow 20FC0404.PV FCCGSU heart cut draw flow The following linear equation is used: P = Ax1 + Bx2 + Cx3 + Dx4 + Bias where: P = HDSFD_SULFUR.PV A = Coefficient 1.097890 x1 = GSUFD_SULFUR.PV B = Coefficient –272.28299 x2 = 20PI0802.PV C = Coefficient 7.0273 x3 = 20FC0306.PV D = Coefficient 3.291770 x4 = 20FC0404.PV Bias = 598.81 Stabilizer-bottom sulfur. This model used several inputs:

Process inputs used Tag name Tag description HDSFD_SULFUR.PV HCN sulfur (HDS feed sulfur IPPP estimation) TABLE 3. Process monitoring points used to estimate sulfur level for the FCC feed inlet Description

Tag name

OHCU bottom from tank

79FC803.PV

LS VGO from tank

79FC802.PV

BH VGO from tank

79FC801.PV

OHCU bottom hot feed

7FC6701.PV

HOT feed from AVU

12FIC100.PV

DHDS VGO ﬂow

2FC0708.PV

AVU crude select tag*

crude_select.op

*The sulfur quantity for each of the ﬂow was operator entry. AVU crude select tag is a digital tag pulled from the AVU having three values.

Tag value

Crude type

Bombay High

4,000

High Sulfur

30,000

Nigerian

6,000

Description

0.875

LS VGO from tank

0.9

BH VGO from tank

0.9

OHCU bottom hot feed

0.875

HOT feed from AVU

0.9

AVU crude select tag Select 165 at www.HydrocarbonProcessing.com/RS

Densities

OHCU bottom from tank

DHDS VGO ﬂow

Sulfur quantity, ppm

Process Control and Information Systems 307TI0642.PV

HDS reactor 1st bed inlet temperature 307TI0630.PV HDS reactor 2nd bed bottom temperature 307TI1014.PV Stabilizer bottom temperature. To estimate the sulfur content of HDS feed, the following linear equation is used: P = Ax1 + Bx2 + Cx3 + Dx4 + Bias where: P = STABBTM_SULFUR.PV A = Coefficient 0.115679 x1 = HDSFD_SULFUR.PV B = Coefficient –3.90 x2 = 307TI0642.PV C = Coefficient –3.59673 x3 = 307TI0630.PV D = Coefficient –0.341067 x4 = 307TI1014.PV Bias = 1067.5 TABLE 4. Economic beneﬁt and octane conservation possible through APC Economic function name

Maximization of sulfur

Speed factor

From FIG. 4, the quality estimation using the IPPP has good agreement with the actual sulfur content as measured from unit and lab analyzers. TABLE 4 summarizes the economic functions and RON improvement possible with APC.

PROJECT MILESTONES Implementing APC on the HDS unit has yielded substantial tangible and intangible benefits. While the annual monetary gain is of the order of Rs. 39 lakhs, significant improvement via process control and optimization was achieved as measured through tighter control of the SHU and HDS reactor inlet temperatures. More accurate estimation of the stabilizer-bottom sulfur inferential was possible, which facilitated proper control action via the APC. With tighter control and action via APC, adjusting and preferentially lowering the reactor-inlet temperatures were possible. The effect of crude changes in the atmospheric and vacuum distillation unit is also incorporated into the model. The resultant sulfur changes in the FCC feed are transmitted via means of intermediate calculations and inferential estimations to the final stabilizer-bottom sulfur prediction. Operators now have more confidence when implementing control and optimization strategies. This has resulted in better operations of the refinery. Accordingly, APC was successfully implemented and is yielding expected benefits.

0.10

Economic coefficients

LITERATURE CITED Perry, R. H., Chemical Engineers Handbook, Sixth Ed., New York, McGraw Hill, 1984. 2 Levenspiel, O., Chemical Reaction Engineering, Third Ed., Singapore, John Wiley and Sons, 1999. 3 Stephanopoulos, G., Chemical Process Control, Dorling Kindersley (India) Pvt. Ltd., 2007. 1

MAX_AI

STEAMMIN

MAX_SULFUR

MINFG

100

MINRIT1

MINRIT2

RON improvement RON improvement after MVPC implementation from the rundown stream (MS) of HDS unit

0.114

1 unit of RON improvement corresponds to (1 metric ton of MS processed)

Rs. 91.30

Annual processing of feed (MS) in the HDS unit (not considering the heart cut drawn from FCCU-GS)

376,487 metric ton

Estimated annual beneﬁt due to MVPC application in HDS unit

Rs. 39,32,517.86 ≈ Rs.39. 32 Lakhs (Rupees thirty nine lakhs thirty two thousand ﬁve hundred and seventeen only)

Sulfur in the stabilizer bottom MS stream improved

15 ppmw

Sulfur in the rundown MS improved

11 ppmw

Targeted beneﬁts due to RON improvement Targeted annual beneﬁt due to MVPC application in HDS unit Targeted sulfur improvement in the rundown MS

Rs. 24.46 Lakhs 10 ppmw

SHYAMAL DEBNATH is the chief technical services manager at Indian Oil Corp. (IOC) Ltd.’s Mathura refinery. His primarily responsibilities include providing technical services for strategic initiatives and advanced process control (APC). Mr. Debnath has more than 25 years of experience in unit operations, strategic initiatives (process and projects), research, troubleshooting and APC for all the major process units at various IOC refineries. He holds an MS degree in chemical engineering from Indian Institute of Technology, Kharagpur, India. HITESH SHAH is a senior technical services manager with Indian Oil Corp. (IOC) Ltd.’s Mathura Refinery. His primary responsibilities include providing technical services for strategic initiatives and APC. Mr. Shah has more than 14 years of experience in strategic initiatives, planning and coordination, and APC. At present, he is working as a senior technical services manager at IOC’s Gujarat refinery. Mr. Shah holds an MS degree in chemical engineering from Indian Institute of Technology, Bombay, India. PRASHAT DUBE is a senior process engineer at Indian Oil Corp. (IOC) Ltd.’s Mathura Refinery. He is primarily responsible for providing technical services for APC implementation and maintenance. Mr. Dube has five years of experience in APC for all major process units at the Mathura Refinery and holds a BS degree in chemical engineering from Indian Institute of Technology, New Delhi, India. MS. VARSHA YADAV is a senior process engineer at Indian Oil Corp. (IOC) Ltd.’s Mathura refinery. She is primarily responsible for providing technical services for APC implementation and maintenance. Ms. Yadav has three years of experience in APC for all major process units at the Mathura Refinery and holds a BS degree in chemical engineering from Regional Institute of Technology, Raipur, India. Hydrocarbon Processing | OCTOBER 2012 59

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www.ametekpi.com Select 96 at www.HydrocarbonProcessing.com/RS

Refining Developments P. K. NICCUM, KBR Inc., Houston, Texas

Maximize diesel production in an FCC-centered refinery, Part 2 Part 1 of this article, published in September, presented several methodologies for maximizing the production of high-quality diesel in a refinery that relies on fluid catalytic cracking (FCC) as its principal means of heavy oil conversion. Part 2 focuses on the selection of FCC catalysts, methods for hydroprocessing light cycle oil (LCO) from the FCC unit, and the production of diesel fuel from FCC byproducts, among other topics. FCC catalyst selection. Some catalyst recommendations apply to both high-severity and low-severity FCC operations. Low-hydrogen-transfer FCC catalyst is recommended for maximizing refinery diesel production, as this type of catalyst will generally produce a higher-yield and higher-quality LCO that can be hydroprocessed, while increasing the yield of FCC olefins that can be oligomerized. Similarly, active matrix functionality improves LCO yield and quality. H2 transfer reactions strip H2 from saturated LCO molecules (such as naphthenes) and transfer it into gasoline boilingrange olefins. The net impact of these H2 transfer reactions is that the LCO becomes more aromatic (lower cetane number and more dense), the gasoline becomes more saturated (lower olefin content and lower octane), naphtha yield increases, and LPG olefin yield declines. In FCC operations intended to maximize gasoline production, the H2 transfer reactions provide a net benefit due to the increased gasoline volume resulting from the saturation of the gasoline olefins before they catalytically crack into LPG olefins. The negative impact of H2 transfer activity on LPG olefins, and on naphtha yields and naphtha octane, has been widely documented, while the negative impact on LCO quality has been less publicized. In high-LCO-yield FCC operations where LCO quality, gasoline octane and LPG yield considerations are more important than sheer gasoline volume, H2 transfer reactions are counter-productive. Refer to TABLE 1 for an example of how the rare-earth content of FCC catalyst can impact FCC yields and product qualities.1 The base catalyst can also be used in combination with a ZSM-5-containing catalyst additive to further preserve the gasoline octane and C3/C4 olefins at low conversion levels. The ZSM-5 additive is applicable to maximizing olefins production from high-severity FCC operations.1, 5 The data in TABLE 2 provide an example of how a ZSM-5 additive can change the yields and product qualities in a moderate-severity FCC operation.2

In low-severity, high-LCO-yield FCC unit operations, ZSM-5 additives have also been shown to convert higher-boiling FCC products into both gasoline and LPG. Two examples of the impact of ZSM-5 additions in low-severity FCC operations are shown in TABLE 3. These data show that the cracking of heavier molecules in the low-severity FCC products by the ZSM-5 results in a loss of total cycle oil (302°F–698°F) production, along with increases in both 302°F true-boiling-point (TBP) gasoline and LPG production.3 Based on a large sampling of pilot plant product from runs having an average conversion level of 40% and a 0.5-wt% ZSM-5 crystal addition, the average Research Octane Number (RON) changes were as follows: • Increase of 2.4 numbers for the initial boiling point (IBP) to 302°F gasoline • Increase of 3.3 numbers for the IBP to 410°F gasoline. Low-equilibrium catalyst micro-activity testing (MAT) activity is often employed when maximizing LCO production. Active-matrix FCC catalysts are also recommended for LCO TABLE 1. FCC pilot plant comparison of yields and product qualities with different catalysts* Catalyst Conversion, vol%

Higher-rare-earth REY catalyst

Lower-rare-earth USY octane catalyst

72.5

0.02

1.28

1.13

1.9

1.4

Yields H2, wt% C1 + C2, wt% C3, vol% C3 =, vol%

7.6

C4s, vol%

13.6

15.1

Gasoline, vol%

LCO, vol%

18.1

19.5

640°F residue, vol%

9.4

Coke, wt%

4.6

90.4

Gasoline octane, RON + 0 Gasoline octane, MON + 0 LCO gravity, °API LCO aniline point, °F

18.4

20.1

*Constant pilot plant feedstock and operating conditions: 23.9°API VGO, 40 weight hourly space velocity (WHSV), 4 catalyst:oil weight ratio (C/O), and temperature of 950°F.

Hydrocarbon Processing | OCTOBER 2012 61

Refining Developments maximization, as they enable the cracking of LCO boiling-range aliphatic side chains from high-molecular-weight feed components. In addition to increasing LCO yield, the aliphatic side chains that report to the LCO boiling range improve LCO cetane. The active matrix also contributes to cetane improvements because matrix cracking does not possess the higher H2 transfer characteristic of a zeolite. Refer to TABLE 4 for representative data concerning the impact of changing the catalyst matrix activity.4 Maximize LCO endpoint. The maximization of LCO endpoint is a common operating strategy for increasing LCO production at the expense of low-value FCC slurry oil. In many FCC operations, concern for coking in the FCC main fractionator bottoms circuit limits the LCO endpoint. A number of FCC operating parameters influence the propensity of the bottoms circuit to suffer coking problems: TABLE 2. Effect of ZSM-5 additive on yields and product qualities in FCC pilot plant* Catalyst

Octane-barrel Catalyst with 4% FCC catalyst ZSM-5 additive

Conversion, vol%

Delta

2.38

2.49

0.11

Yields H2, C1+ C2, wt%

1.9

−0.1

C3=, vol%

C3, vol%

6.8

7.4

0.6

C4=, vol%

6.1

6.9

0.8

iC4, vol%

4.2

−0.2

nC4, vol%

1.1

0.9

−0.2

Total LPG

20.2

21.1

0.9

Gasoline (450°F TBP), vol%

57.6

56.8

−0.8

LCO, vol%

17.9

−0.1

Bottoms, wt%

14.1

0.1

3.9

3.8

−0.1

Gasoline octane, RON + 0

Coke, wt%

90.2

91.6

1.4

Gasoline cetane, MON + 0

79.2

79.6

0.4

*Constant pilot plant feedstock (27.0°API VGO) and operating conditions (960°F).

TABLE 3. FCC plant data showing effect of ZSM-5 additive on yields and product qualities Low-conversion FCC operation Catalyst system

FCC product considerations. Changes in FCC cracking

severity directly impact FCC product yield distribution and qualities. In the FCC pilot plant example presented in TABLE 5, the VGO is of average quality as an FCC feedstock, and the catalyst is a low-rare-earth catalyst with some matrix activity. The pilot plant runs covered reactor temperatures and conversion levels ranging from low to high, relative to industry norms. The pilot plant data show the tradeoffs between LCO production and quality, and the production and quality of FCC naphtha. As shown in TABLE 5, even without adjusting the LCO cutpoints, the LCO yield changes by a factor of almost 2 by adjusting the FCC reaction severity. At the same time, among the runs presented in TABLE 5, the gravity of the LCO increases by about 11°API as the operating severity is lowered. FIG. 1 summarizes the positive relationship between increasing LCO production rate and LCO quality, as observed in a TABLE 4. FCC pilot plant study results Catalyst matrix surface area

Low

High

Conversion

69.5

69.7

53.1

Gasoline (C5 at 421°F) Yield RON

87.7

MON

77.8

78.5

36/23/15/27

26/36/14/24

Paraffins/oleﬁns/napthenes/ aromatics (PONA) LCO (421°F–602°F) Yield

16.3

19.2

Cetane index

24.5

28.5

API

21.8

23.8

Aromatic carbon, %

49.5

45.9

Aliphatic carbon, %

50.5

54.1

Carbon NMR

Plant A

Plant B

REY zeolite with ZSM-5 additive

Incremental yields from ZSM-5 addition Dry gas, wt%

• Bottoms circuit temperature • Bottoms circuit liquid residence time • Concentration of unconverted paraffins in the slurry oil. In high-conversion FCC operations, the slurry oil is more aromatic and can be held at higher temperatures and longer residence times without coking. Some of the slurry oil quality data that FCC operators monitor as indicators of coking tendency are gravity and viscosity. The more aromatic slurry oil produced by high-conversion FCC operations will allow the unit to operate with lower API gravities while respecting bottoms viscosity targets selected to avoid fractionator coking.

Bottoms (602°F+) Yield

14.2

11.1

7.6

+0.3

–

Gravity, °API

LPG, vol%

+2.4

+2.9

Carbon NMR

Gasoline (302°F IBP), vol%

+4.8

+3.3

Aromatic carbon, %

57.4

Total cycle oil (302°F–698°F)

–3.2

–6.7

Aliphatic carbon, %

42.6

Bottoms (698°F+)

–4.5

+0.2

Viscosity at 210°F, cst

7.87

5.8

–

+0.2

Viscosity at 100°F, cst

116.4

68.14

Coke, wt%

62 OCTOBER 2012 | HydrocarbonProcessing.com

Refining Developments larger sampling of the same pilot plant study data. Conversely, FIG. 2 and FIG. 3 show a very direct and negative correlation between LCO yield and FCC naphtha octane. FIG. 2 demonstrates that, irrespective of the indicated FCC reaction temperature, FCC naphtha motor octane will suffer as LCO yield increases. FIG. 3 shows that the negative impact of increasing LCO yield on the olefin-dependent RON can be mitigated to some extent, if a high FCC reaction temperature is maintained. The data in TABLE 5 also provide examples of how changing FCC reaction severity can impact LPG yield and naphtha octane. Comparing the low-conversion and high-conversion cases, the data show that the low-conversion case produces less than one-half the LPG and 3 to 4 numbers lower octane than the high-conversion case. TABLE 5 also provides an example of the degradation of LCO as a potential feedstock for upgrading into diesel as the FCC

conversion is increased; the LCO H2 content decreases from 10.7 wt% to 8.8 wt% as the FCC conversion level is increased from 59 wt% to over 76 wt%. Hydroprocessing options. Processes for the upgrading of LCO range from mild hydrodesulfurization to full-conversion hydrocracking. FIG. 4 depicts some of the chemistry responsible for improving the cetane, density and aromatics content of the LCO. For the purposes of this article, three upgrading processes (hydrotreating, aromatics saturation and mild hydrocracking) are described as representative examples of some of the processes being used today.5 LCO hydrotreating. Mild hydrotreating of LCO will reduce its sulfur content significantly, but this will only modestly improve the product qualities related to aromatic content. In examples presented in TABLE 6, LCO in a 10% concentration, in

TABLE 5. FCC pilot plant data showing impact of changing operating severity Low conversion

Medium conversion

High conversion

FCC feed properties Gravity, °API

22.5

50 vol% boiling point, °F

851

Aniline point, °F

176

Sulfur, wt%

0.55

CCR, wt%

0.89

FCC pilot plant operating conditions Riser temperature, °F

940

979

1,020

Feed temperature, °F

416

485

337

Catalyst-to-oil ratio, wt/wt

6.6

6.7

11.4

Micro Activity Test (MAT)

0.6

Dry gas, wt%

1.23

2.08

3.5

C3 LPG, wt%

2.97

4.26

7.27

C4 LPG, wt%

5.98

7.88

11.57

Gasoline (C5 at 430°F), wt%

43.21

46.98

LCO (430°F–680°F), wt%

27.42

24.47

16.01

Slurry oil (680°F+), wt%

13.6

9.06

7.66

Coke, wt%

5.59

5.27

7.99

58.98

66.47

76.33

Rare-earth oxides, wt% (FCC E-Cat property) FCC pilot plant yields

Conversion, wt% FCC pilot plant product qualities C3 LPG oleﬁnicity, wt%

83.8

85.7

C4 LPG oleﬁnicity, wt%

66.7

68.5

Naphtha gravity, °API Naphtha octane, RON/MON Naphtha PONA, wt%

56.6

57.2

55.9

91.7/81.1

92.9/81.6

95.6/84.4

27.2/49.5/11.8/11.5

25.7/49.1/10.9/14.3

31.3/36.8/10.5/21.4

LCO gravity, °API

22.2

11.3

LCO H2 content, wt%

10.7

9.9

8.8

Slurry oil gravity, °API

−0.8

−7.4

Slurry oil H2 content, wt%

7.8

6.7

Hydrocarbon Processing | OCTOBER 2012 63

Refining Developments a mixture including straight-run gas oil (SRGO), is hydrotreated. Two options are presented, with the latter representing a higher degree of desulfurization and greater aromatics reduction. These examples demonstrate that it is possible to include about 10% LCO in the diesel pool by hydrotreating the LCO/ SRGO mixture. Aromatics saturation. To accommodate larger concentrations of LCO in the diesel pool, more complete aromatics LCO quality

saturation and cetane improvement are required. These goals can be achieved through varying degrees of ring saturation and ring opening, as shown in FIG. 4. TABLE 7 shows what is possible utilizing a two-stage aromatics saturation unit to process 100% LCO.5 The drawback of ring saturation is high H2 consumption. Mild hydrocracking. Another alternative is to rely on ring opening with mild hydrocracking to move some of the aromatics out of the LCO boiling range into gasoline, as shown TABLE 6. Processing a 10% LCO blend with ULSD catalyst systems

940°F 980°F 1,020°F

Product

Gravity, °API

Operating pressure 20

Feed: 90% SRGO/10% LCO

CoMo

NiMo

Medium

High

880

863

853

15,300

543

534

523

Density, kg/m3 Sulfur, wppm

D86 T10, °F

10 5 10

586

579

570

D86 T90, °F

660

657

649

IP391 monoaromatics, wt%

16.7

22.6

21.4

9.2

2.8

31.8

24.2

IP391 PNA, wt%

LCO yield, wt%

D86 T50, °F

IP391 total aromatics, wt% FIG. 1. Relationship between increasing LCO production rate and LCO quality.

52.5

H2 consumption, Nm3/m3

FCC naphtha quality

TABLE 7. Two-stage LCO aromatics saturation 940°F 980°F 1,020°F

84 Naphtha MON

Cetane number

100% LCO

Two-stage

Operating mode

Product

Medium

Operating pressure 960

859

7,300

< 10

79.1

2.5

Density, kg/m3

Sulfur, wppm 81

Total aromatics (FIA), vol%

24.1

40.2

< 20

44.9

25+

Liquid yield, vol%

115.7

H2 consumption, Nm3/m3

473

Cetane index, D976

80 10

Cetane number

LCO yield, wt%

Delta cetane number FIG. 2. Relationship between FCC naphtha quality (MON) and LCO yield.

FCC naphtha quality

940°F 980°F 1,020°F

Naphtha RON

Aromatic saturation

1 2H2

Diesel

3H2

94 93

C5H11

91 90 10

LCO yield, wt%

FIG. 3. Relationship between FCC naphtha quality (RON) and LCO yield.

64 OCTOBER 2012 | HydrocarbonProcessing.com

Diesel Selective ring opening

3H2 3

FIG. 4. Three reactions to upgrade LCO quality.

C5H11

Hydrocracking Diesel

Refining Developments TABLE 8. ULSD and mild hydrocracking on feed blend containing 10% LCO and 35% coker diesel* Property

Feed

ULSD product

MHC product

Density, kg/m3

866

842

829

822

Delta density

Sulfur, wppm

8,000

< 10

42.4

13.2

Mono

12.8

13.5

PNA

12.4

0.4

0.5

36.8

43.8

46.2

46.8

SFC aromatics (total), wt%

Total product cetane index, D4737 Delta cetane index

9.4

Chemical H2 consumption, Nm3/m3

116

150

155

Incremental 379°F minus, vol%

1.1

*For MHC cases, diesel product is 2 to 3 cetane numbers higher than total product.

100 80 Cetane number

LPG Naphtha

N-paraffin Mononaphthenes

Selective ring opening

40 20

Aromatic naphthene

Diaromatics

-20 100

150

Aromatic saturation

FCC C3/C4 LPG

Monoaromatics

Diesel

HDS H2

Dinaphthenes

H2S

FCC recycle

Raffinate: paraffins + oleﬁns FCC naphtha

200 Molecular weight

Oleﬁn oligomerization unit

LCN

250

Extract: sulfur + aromatics

MCN

FIG. 5. Hydrocarbon comoponents and cetane number.

in FIG. 5. This approach can provide substantive LCO quality improvement with lower H2 consumption. TABLE 8 provides an example of coprocessing LCO along with straight-run distillate and other cracked products.5 Creating diesel from FCC byproducts. Two processing op-

tions with limited application to date are the creation of synthetic diesel from FCC olefins and the extraction of aromatics from FCC naphtha. These options can be integrated into the overall processing scheme, along with the other options described earlier. Reprocessing of C3–C9 olefins into distillate. Olefins can be used to produce good-quality diesel with oligomerization processes. For example, an oligomerization unit distillate yield from a C3–C9 olefin feed was reported to be 78% distillate with a byproduct gasoline yield of 19%, based on a zeolite catalyst, as shown in TABLE 9. After hydrotreating to saturate the olefins, the distillate was reported to have a cetane number of 52 to 54, zero sulfur and less than 2% aromatics.6 Therefore, for FCC-based refineries working to maximize diesel production, oligomerization of olefins-containing FCC light gasoline and LPG may provide viable investment opportunities. FCC naphtha extraction. Extractive techniques are available for separating a middle boiling fraction of FCC gasoline into a higher-octane, aromatics-rich fraction and an olefinsand paraffin-rich fraction.7 A recently granted patent describes a combined FCC/extraction process wherein an aromatics-

H2S

Solvent HCN

Aromatics

HDS

300

Severe HDS

ULS gasoline blending

FIG. 6. Production of diesel from FCC LPG and FCC naphtha.

rich, higher-octane fraction of FCC gasoline can be produced as a gasoline product, while a paraffinic/olefinic naphtha fraction can be produced for recycle to an FCC riser for the purpose of producing propylene and other olefins.8 This FCC naphtha extraction concept and oligomerization technology can be used together, as shown in FIG. 6, to maximize the production of synthetic diesel from FCC olefins. The combination can be especially useful in the context of a high-LCO-yield, low-severity FCC operation because the lowseverity FCC naphtha will have a higher olefins content than the more aromatic, more paraffinic naphtha from a high-severity FCC operation. Thus, the non-aromatic naphtha raffinate from a low-severity FCC operation will make a better-quality oligomerization feedstock—or a better-quality FCC recycle stream—for the purpose of increasing lighter FCC olefins production, as olefins are easier to crack than paraffins. Refinery diesel balance. With all the processing options presented in this article, an obvious question is, “How much can the refinery diesel production be increased if many of these options are applied in a retrofit of an existing refinery?” The answer depends on the specifics of the application. TABLE 10 shows estimated results from isolated examples provided in this article, giving insight into the question. Hydrocarbon Processing | OCTOBER 2012 65

Refining Developments Takeaway. Assuming demand for diesel continues to increase faster than growth in gasoline, a number of reactions can be expected from the refining industry: • The loss of virgin diesel to the FCC unit will diminish through crude distillation unit improvements • FCC gasoline endpoint will be minimized • Hydrocracking and hydrotreating units designed to upgrade LCO quality will proliferate • Low-H2 -transfer, higher-matrix-surface-area FCC catalyst will be used to improve LCO yield and quality, while increasing LPG olefins production and naphtha octane • In some cases, ZSM-5 catalyst additives will be used to further increase LPG olefins production and octane, but in low-severity FCC operations, this may come at the expense of some LCO yield. For refiners that also place high value on propylene production, high-octane gasoline, and minimization of refinery bottoms production, the high-severity FCC route to making more diesel will gain favor through the oligomerization of C4 and higher FCC olefins while continuing to hydroprocess the LCO production. If a refiner has a more singular focus on the production of diesel, the low-severity, traditional FCC route to increasing diesel can be optimized and economically favored, with some enhancements: • The loss of LCO in slurry oil product or recycle will diminish through the use of dedicated slurry distillation hardware TABLE 9. Product yields and properties from oligomerization of oleﬁns Feed composition

82% C3–C9 oleﬁns

Product yields (based on feed oleﬁns), vol% Gasoline

Distillate

• Some of the stripped slurry oil may be recycled to the FCC reactor to produce more LCO and help maintain FCC heat balance, while HCO recycle may also be advantageous • Low-severity FCC operations will rely on increasing feed temperature and, in some cases, direct firing of the regenerator with a liquid or gaseous fuel using technology designed to minimize damage to the catalyst • FCC-produced LPG and naphtha olefins will be converted into diesel blending stock using oligomerization processes. An ultimate vision for maximizing diesel production in a specific FCC-centered refinery may also include a selective combination of elements: • Extraction processes will separate aromatics-rich fractions of FCC gasoline from fractions enriched in olefins and paraffins. The aromatic fraction can be used for BTX production or high-octane motor fuel; the non-aromatic fraction can be recycled to the FCC reactor for the production of more olefins (diesel precursors), or the olefins in the non-aromatic fraction may be directly oligomerized into diesel. • FCC C4s and FCC light naphtha can be recycled to an ultra-high-severity FCC riser to increase propylene and aromatic naphtha yields, without diminishing LCO production. A case-by-case analysis based on refinery-specific data is needed to accurately contrast the costs and benefits associated with the application of various options for increasing diesel production from the FCC-centered refinery. The performance of the study requires both refinery-wide and FCC-specific experience and related modeling capabilities. In the final analysis, it is simply a question of economics; technologies are available to maximize diesel production from the FCC-centered refinery. LITERATURE CITED Complete literature cited available online at HydrocarbonProcessing.com.

Distillate qualities after mild hydrotreating Boiling range, °F (IP 123/84)

388–676

Density, kg/m3 at 20°C Cetane number

787 52 to 54

Aromatics content, wt%

Sulfur content, wt% Viscosity, cst at 104°F

0 2.55

PHILLIP NICCUM joined KBR Inc.’s fluid catalytic cracking (FCC) team in 1989, following nine years of FCC-related work for a major oil company. Since that time, he has held various FCCrelated positions at KBR Inc., including process manager, technology manager, chief technology engineer of FCC, director of FCC technology, and now process engineering manager. Mr. Niccum’s professional activities have included engineering management, process engineering, project engineering, marketing, and licensing. Areas of technical strength include FCC unit design, precommissioning and startup, troubleshooting and economic optimization of FCC unit operations.

TABLE 10. Examples of reﬁnery diesel increases Vol% of reﬁnery crude input (assuming 27% crude oil to FCC unit and LCO hydroprocessing to maximize diesel in all cases)

Low-severity FCC

Moderate-severity FCC

High-severity FCC

Change FCC severity (assuming constant MAT and no recycle)

0.9

–

–2.5

Minimize diesel in FCC feedstock

3.2

3.4

3.6

Lower FCC naphtha endpoint

5.1

5.3

4.6

Change FCC catalyst formulation (increase FCC catalyst matrix activity and reduce H2 transfer activity)

1.3

Refractionate slurry oil (recover 30 vol% LCO from slurry)

1.1

0.7

0.6

Oligomerize C3–C5 oleﬁns

3.1

3.7

5.1

Oligomerize C6+ naphtha oleﬁns

3.8

4.3

2.6

18.4

18.6

15.3

Total increase in diesel production, vol%

66 OCTOBER 2012 | HydrocarbonProcessing.com

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SULFUR

OPTIMIZE SULFUR RECOVERY FROM DILUTE H2S SOURCES M. P. HEISEL, ITS Reaktortechnik GmbH, Pullack, Germany; and A. F. SLAVENS, WorleyParsons, Monrovia, California

With the recent emergence of “sweeter” fuel sources such as unconventional gas and biofuels, the sulfur industry is challenged in providing economical solutions to desulfurize gas streams with low hydrogen sulfide (H2S) content. Historically, the sweetening of such gases was primarily accomplished by using liquid redox processes.a These established, sweetening processes treated the raw gas stream directly and did not require an amine treating step, thereby reducing the total capital and operating costs of the facility. Liquid redox processes are capable of scrubbing the H2S to very low levels and meeting typical treated-gas specifications, as proven by several hundred units that are in operation. However, these facilities often suffer from very high operating costs, low availability and a lowquality sulfur product, which usually must be disposed rather than sold as product. The inherent nature of these problems is discussed here: • High operating costs are a result of the process chemistry, especially consumption of the expensive chelating agents required to keep the direct oxidation catalyst in solution. Consequently, chemical costs range from $100 to $150 per ton of produced sulfur. • The product sulfur contains some chelating agent and, therefore, is a low-quality material. As a consequence, no revenue from sulfur sales can be expected. More important, additional costs associated with landfill disposal can be incurred. • Low process availability results from two primary steps, as shown in FIG. 1. In the scrubber (1), sour gas is contacted by liquid solvent and thus forming the solid sulfur, which often leads to plugging in the column or in downstream vessels and pipes. In the re-oxidation vessel (5), foaming and sulfur froth can occur, thus reducing availability.

New technology. A newly developed process applies a totally different approach.b This process first oxidizes H2S selectively in the gas phase over a robust and low-cost catalyst. To increase sulfur recovery efficiency above what was achievable in the selective oxidation step, the process subsequently applies the sub-dewpoint principle. It is a wellknown Claus tail-gas treatment technology that takes advantage of the improved Claus equilibrium at lower operating temperatures (below the sulfur dewpoint) in the catalytic reactors. The process can achieve sulfur recovery efficiencies exceeding 99% when treating in low H2S-content gases, such as shale gas, coalbed methane and biogas. The process is inexpensive and easy to operate; it generates no byproducts, and the sulfur recovered is of premium quality. The direct oxidation process is capable of treating raw gas streams containing H2S plus hydrogen, light hydrocarbons, oxygen and/or inert gases. FIG. 2 shows a flow diagram of the new direct oxidation process.

Process description. The feed gas to the sulfur recovery unit (SRU) is mixed with a stoichiometric quantity of air to convert the incoming H2S to elemental sulfur via direct oxidation. The gas mixture is sent through a preheater to the first reactor. This reactor is different from conventional Claus reactors: it contains two sections. The upper section at

the gas inlet is a conventional fixed-bed reactor with a direct-oxidation catalyst. In this reactor section, part of the feed H2S is oxidized into elemental sulfur according to Eq. 1. In parallel, some sulfur dioxide (SO2) is formed. The second section in the lower part of the first reactor contains a Claus catalyst with an embedded heat exchanger, which is designed to remove the heat of reaction from the catalyst bed. The heat removal within the catalyst bed shifts the equilibrium of the Claus reaction (Eq. 2) toward more sulfur formation, substantially improving conversion efficiency. Direct oxidation of H2S 2 H2S + O2 = 2/x Sx + 2 H2O + heat of reaction

(1)

Puriﬁed gas Vented air

Catalytic solution

Sour gas

LP ﬂash gas

4 Sulfur Air

FIG 1. Typical process flow diagram of a liquid redox process.1 Air

Air blower 4-way valve

Feed gas Preheat

Reheat Selective oxidation reactor

Recycle blower (optional)

Puriﬁed gas

4-way valve Sulfur condenser

Sub-dewpoint reactor

Sulfur separator

Sulfur pit C01 Air blower, E01 Preheat, V01A/B 4-way valves, R01 A/B Reactors with internal cooling, E02 Sulfur condenser, D01 Sulfur separator, E03 Reheat, P01 Sulfur product pump, C02 Recycle blower (optional)

FIG. 2. Typical process flow diagram of new direct oxidation process for H2S. HYDROCARBON PROCESSING

SULFUR 2012

S-69

SULFUR Claus reaction 2 H2S + SO2 = 3/x Sx + H2O + heat of reaction

(2)

where x = 2, 4, 5, 6, 8 sulfur molecules of different sizes, according to temperature. The heat exchanger applied is a thermoplate stack with large clearances, as shown in FIG. 3. The space between the thermoplates is filled with catalyst. As this heat exchanger type is not yet so well known within the sulfur industry, it will be discussed in more detail later. A sulfur condenser is located downstream of the first reactor. A second reactor, identical to the first reactor, follows the sulfur condenser but operates at lower temperature. This shifts the chemical Claus equilibrium to even more sulfur formation. The reactor outlet temperatures range from 100°C to 125°C, i.e., possibly even below the sulfur solidification point. When operating below the sulfur dewpoint, the sulfur formed via the Claus reaction accumulates on the catalyst. Thus, the catalyst deactivates slowly and must be regenerated. The regeneration is accomplished by switching the second reactor into the first reactor position. In the first reactor position, the inlet temperature approaches 320°C, which desorbs sulfur and regenerates the catalyst. The former first reactor is switched at the same time into the second, cooler reactor position. This procedure is repeated typically once every 24 hours. Treated gas from the second reactor is sent to the consumer, e.g., as purified biogas or natural gas.

Process capabilities. The new direct oxidation process can be applied to a number of plants such as for: • Biogas purification • Offgas treatment from chemical processes rich in methane, carbon dioxide and hydrogen • Natural gas purification. Pure, bright yellow elemental sulfur is produced. The process operation is fully automatic, with manual control only required during startup and shutdown, similar to a conventional Claus plant. The first commercial unit was installed in 1993 and is still in operation. The high sulfur-recovery rate (SRR) in the new oxidation process results from removing the heat of reaction, which shifts the chemical equilibrium to more product formation. FIG. 4 illustrates the effect on SRR, where the SRR is depicted as a function of the outlet temperature from the second reactor. In addition to high-sulfur recovery efficiency, the internally cooled reactors provide other benefits. The internal heat exchangers are selfcontrolling. Boiler feedwater (BFW) feeds the inside of the thermoplate; the BFW always has a temperature corresponding to the generated steam pressure. On the outside, i.e., between the thermoplates, the catalyst and reaction gas release heat. The greater the gas flow, the higher the heat of reaction, which is the temperature difference between the gas and BFW. With the higher temperature ⌬, the heat of reaction is automatically removed by internal cooling. As a consequence, the temperature at the outlet of the reactors is constant within a narrow range, and is independent of fluctuations in gas volume and gas composition. Accordingly, this process is intrinsically stable, easy to operate and has high reliability. Actually, all normal operations are fully automatic, thus very little operator attention is necessary. Commercial unit. The first commercial plant applied a two-reactor process configuration; it was started up in December 1995 in the Nynäs refinery in Sweden (FIG. 5). This unit processes amine-acid gas and sour-water-stripper gas. The plant has proven to be very reliable, easy

FIG 3. Top view of a thermoplate heat exchanger for a reactor during construction in the shop. Solid sulfur

100.0

Liquid sulfur

99.9 99.8 SRR, %

99.7 99.6 99.5 99.4 99.3 99.2 100

105

110

115

120 125 Temperature, °C

130

135

FIG. 4. Sulfur recovery rate as a function of outlet temperature of the second reactor for rich feed gas with 85% H2S. S-70

SULFUR 2012 HydrocarbonProcessing.com

140

FIG. 5. Sulfur recovery plant in the Nynäs refinery in Sweden.

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to operate, and requires very little maintenance. Availability is always better than 99.5%/yr. The refiner claims that this plant is the most reliable within the whole refinery, even after more than 15 years of operation. It achieved the required SRR and reached optimal values of up to 99.85%, even with aged catalysts. In operation at low load conditions (at 6:1 turndown), the SRR dropped by only 0.1%.

FIG. 6. Schematic of a thermoplate heat exchanger.

Thermoplates for internal cooling of catalytic reactors. Internally cooled catalytic reactors have been successfully used in many applications. They are applied primarily for selective reactions, where rigorous temperature control is required, or in reactions where the chemical equilibrium is strongly temperature dependent. In the past, straight-tube reactors, with the catalyst inside the tubes, have typically been used. In a few cases, spiral-wound tubular heat exchangers have been applied with the tubes submerged in the catalyst. However, these reactor types have features that are not complementary to the operations. Primarily, the heat exchangers’ fabricated geometry often forces conditions on the catalytic reactions that are not optimal. For example, the straight-tube reactors had to be built slim and high to avoid excessive thermal stress on the tube sheets, which resulted in a high pressure drop, high linear gas velocity and mechanical stress on the lower catalyst particles. The spiral-wound heat exchangers avoid these disadvantages to some degree, but they require many manufacturing steps and precise fabrication skill. They are typically more expensive. Both types of reactors cannot be built onsite, and, therefore, one must observe transportation limitations. This also limits throughput capacity. In view of ever-increasing plant sizes, this condition becomes increasingly more important. All of these features for tubular reactors are detrimental for sulfur recovery, which may explain why internally cooled reactors have not been used widely in sulfur recovery previously. The catalytic reactors incorporated in the new-generation direct oxidation process use thermoplates as heat exchangers, thus eliminating all of the listed disadvantages. The basic element of a thermoplate heat exchanger is the thermoplate itself, as shown in FIG. 6.

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A thermoplate consists of two metal sheets welded together along their edges and point-welded across their surfaces. This is accomplished with precise fabrication machinery that facilitates the manufacturing of large surface area exchangers at low cost. The plates are expanded by injecting high-pressure liquid between the metal sheets, which opens channels for the cooling medium, as shown schematically in FIG. 6. The expansion generates the typical cushion shape, and the point and seam welds of the thermoplates are gap free. Multiple thermoplates are combined to form a heat exchanger package, which is then inserted in a shell to complete the heat exchanger. For application in a reactor, the catalyst is poured into the spacing between thermoplates, as shown in FIG. 3. Vertical plane walls are formed by the thermoplates and allow easy filling of the catalyst particles. Several thousand such thermoplate heat exchangers have been built and installed worldwide. They are in service in even the most severe applications, such as condensing phosgene, which is not only highly toxic, but is also very corrosive when in contact with water. Single heat exchangers with several thousands square meters of exchanger surface area have been installed and operated. The thermoplate heat exchanger is considered a proven technology. These exchangers are compact and light weight, have low pressure drop, and provide high heat exchange coefficients; they are ideal for sulfur-recovery reactors. The outer and inner fluid channels are completely separated from each other by seam welds. As in contrast to other plate-heat exchangers, there is no contact between adjacent thermoplates; each thermoplate is self-contained and no forces are transferred to the next plate. The catalyst particles are insulated and do not experience mechanical stress. The distance between plates, file height, pitch of the point welds,

dimensions and number of thermoplates can vary widely. Therefore, thermoplate reactors can be optimally tailored to each sulfur recovery application.

Options. The new direct oxidation process is an economic method for sulfur recovery from low-H2S content gases. It converts H2S in a gas catalytic process directly to elemental sulfur. The sulfur recovery efficiency, which depends on the feed-gas composition, is greater than 99%. The reaction takes place in two identical fixed-bed reactors with internal cooling by thermoplate heat exchangers, which maintain the outlet temperatures of the reactors within a narrow range, thus maintaining a constant SRR. This process has proven to be easy to operate, very reliable and with low maintenance costs. As one customer commented, “Our biggest problem with this process is that the operators tend to forget about it, because it requires so little of their time and attention.” LITERATURE CITED 1

www.prosernat.com/en/processes/gas-sweetening/sulfint-hp.

Liquid redox processes include LO-CAT, SulFerox and Sulfint. SMARTSULF is a new sulfur oxidation process.

NOTES b

Michael Heisel, PhD, is general manager of ITS Reaktortechnik GmbH. He has more than 30 years of experience in sulfur recovery plant design, startup, validation and troubleshooting.

Angela Slavens is vice president and global director of sulfur technology for Worley Parsons. She has more than 15 years of experience in the oil and gas industry, primarily in the field of sour gas treating and sulfur recovery.

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CB&I

CB&I COVERS THE ENTIRE PROJECT LIFECYCLE, CONCEPT TO COMPLETION From humble beginnings nearly 125 years ago, CB&I has continually expanded its capabilities to serve the energy and natural resource industries. Today, CB&I engineers and constructs some of the world’s largest energy infrastructure projects. With premier process technology, proven EPC expertise and unrivaled storage tank experience, CB&I executes projects from concept to completion. We offer a comprehensive range of capabilities that span the entire project lifecycle: CB&I’s Project Engineering and Construction business sector builds upstream and downstream oil and gas projects, LNG production and regasification terminals, and a wide range of other energy related projects. CB&I’s Steel Plate Structures business sector designs, fabricates and constructs storage tanks and containment vessels and their associated systems for the oil and gas, water and wastewater, mining and nuclear industries. CB&I’s Lummus Technology business sector provides proprietary process technologies, catalysts and specialty equipment to petrochemical facilities, oil refineries and gas processing plants. Safety is a core value at CB&I and we are proud to have one of the best safety records in the industry. Throughout our organization, every employee worldwide is committed to safe work practices. Our awardwinning safety program promotes a culture of involvement and dedication with a goal of zero incidents for everyone involved in our projects.

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CONTACT INFORMATION CB&I 2103 Research Forest Drive The Woodlands, TX 77380 USA Tel: +1 832 513 1000 Fax: +1 832 513 1005 info@cbi.com www.CBI.com

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ENERSUL LIMITED PARTNERSHIP

ENERSUL LIMITED PARTNERSHIP TRUSTED EXPERIENCE. PROVEN EXCELLENCE. Enersul Limited Partnership headquartered in Calgary, Alberta, has been the world leader in the sulphur forming and handling industry for over sixty years. With a complete array of operational, technical, and supportive offerings, Enersul has the unique ability to provide complete sulphur solutions, customizable to fit any production requirement.

OPERATIONAL SOLUTIONS Enersul’s Operational Solutions has a depth of experience unmatched in the field. Long-standing relationships with sulphur producers have established Enersul’s reputation as a leader in reliability, safety, and environmental consideration. Having taken complete operational control of the sulphur requirements of a variety of projects enables Enersul to innovate their offerings to meet the needs of the real world. Lessons learned by Enersul’s international project teams are applied across every aspect of Enersul’s products and services.

TECHNICAL SOLUTIONS From when molten sulphur leaves the SRU to transportation loading for its final destination, Enersul provides customizable technical solutions for every step of the process. Each technology has been designed to meet the strictest safety and environmental standards with a dedicated focus to functional reliability and standard setting end-product quality.

for clients that require smaller production rates. It arrives 90% assembled, with most of the system checks already completed, thus saving on construction and commissioning time and costs. Each unit can produce up to 400 tonnes of high quality granules per day. Enersul’s WetPrill™ product is known for its low friability, lower moisture content and high bulk density as compared to other wet prill products. The WetPrill process units can be scaled for operations ranging from 100 to 2500 tonnes per day with higher throughput achieved with multiple process lines.

SOLID SULPHUR HANDLING AND STORING H2S DEGASSING In response to the need for a compact and efficient sulfur degassing process, Enersul developed the HySpec™ H2S degassing process. These processes has been specifically designed to quickly, effectively and economically reduce the H2S content of liquid sulfur to 10 ppm or less. The in-line, continuous flow design of the HySpec™ process eliminates the need for large molten sulfur pits typically required with traditional batch-type degassing systems. This concept allows for easy retrofitting of existing facilities, reducing the capital cost of degassing system installations. Due to its modular and compact design, a HySpec™ unit can be installed with only minor tie-ins and minimal disruption to ongoing operations.

LIQUID SULPHUR SOLUTIONS

Enersul provides a variety of conveyor systems to safely and efficiently handle formed sulphur designed to reduce end-product degradation. Transfer points are kept to a minimum, drop distances are minimized, and covered or non-covered handling systems are designed to ensure formed sulphur retains the high quality only available from Enersul’s patented forming technologies.

TRUSTED EXPERIENCE. PROVEN EXCELLENCE. Enersul specializes in one thing, the safe, reliable, and environmentally friendly forming and handling of sulphur. Over 60 years of innovation with a focus on end-product quality, and the ability to customize technical and operational solutions to any plant requirement means the worlds sulphur needs can rely on Enersul’s Trusted Experience and Proven Excellence.

Enersul develops systems to pipe, store, filter, cool, and load molten sulphur. Safety, the environment, and reliability are constant considerations applied to flexible executions to meet any specific plant production requirements or shipping schedule.

SULPHUR FORMING The GXM1™ has a forming capacity of 1250 tonnes per day. It is completely self-contained: a compact design with a rotating drum, a cooling water system, a wet scrubber, vibrating screens and conveyors. Operation is simple so on-time performance is high, labor requirements are minimal, and costs for repair, maintenance and utility consumption are kept low. The GXM3™ is the first SinglePass™ granulation technology and the only portable sulfur forming unit on the market. This patent pending technology was developed for use on sites with smaller plant footprints and SPONSORED CONTENT

CONTACT INFORMATION 7210 Blackfoot Trail SE Calgary Alberta Canada T2H 1M5 Phone: (403) 253-5969 Fax: (403) 259-2771 E-Mail: enersul@enersul.com www.enersul.com HYDROCARBON PROCESSING

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What you can do

with a

You can…

touch of blue. Enhance the efﬁciency of your overall sulfur recovery to achieve peak operating and environmental performance with our SRU technology. Lower burner operating temperatures with our proprietary acid gas burner technology, ultimately extending the operating life of the burner and reducing operating and maintenance costs. Maintain environmental compliance with sulfur recovery efﬁciencies up to 99.9+% to meet the most stringent environmental regulations. Replace your burner in an existing plant.

Our high performance technology, coupled with our focus on aftermarket support and training, delivers the result you need. Visit www.fwc.com/touchofblue for more information on our sulfur recovery technology. Select 83 at www.HydrocarbonProcessing.com/RS

FOSTER WHEELER

OPERATING, ENVIRONMENTAL SOLUTIONS WITH SULFUR RECOVERY TECHNOLOGY SULFUR RECOVERY TECHNOLOGY Foster Wheeler’s proprietary and proven sulfur recovery technology brings clear advantages to our refinery customers. The technology provides cost-effective designs with enhanced operability features. Included in the proprietary technology is a Claus unit burner that is capable of destroying ammonia up to 25 mole % in the SRU feed, and providing low level oxygen enrichment up to 28 mole %. The units typically deliver overall sulfur recovery efficiencies ranging from 96% to 99.9+% depending on the configuration of the claus and tail gas treating sections.

OUR EXPERTISE Our personnel employed at the new Foster Wheeler Salt Lake City office are knowledgeable in the design of sulfur block units, including Claus units, tail gas treating units and tail gas incinerators. Our sulfur expertise also includes sour liquid or gas amine absorbers, amine regenerators, sour water strippers, sulfur condensers, waste heat boilers, sulfur storage, sulfur degassing, and sulfur pit vent disposition. Other areas of proficiency include hazardous waste incineration, natural gas processing, general refinery units, and mining and chemical plants.

3 Catalytic Reactor Beds 200LTPD Sulfur Recovery Unit

OUR SCOPE OF WORK Coupled with the small footprint, our design offers reduced piping runs that are completely free draining. Lower corrosion and reduced pressure drop are clear benefits from the reduced pipe routing, which also results in lower Capex and enhanced operability and maintenance. The footprint of our claus units is minimized by combining the waste heat boiler and sulfur condensing tube bundles in a common shell operating at steam pressures matching the refinery steam system, and providing the steam required for the claus unit operation. The mechanical expertise required for reliable and safe design of the waste heat boiler and sulfur condenser tubesheets, as well as the partitioning of the condenser passes in the boiler plenums, has been developed through many years of experience. Steam pressures ranging from 50 psig to 600 psig are available, and each plant is designed to be self-sustaining in steam usage during normal operation. These proven, innovative designs help to set our technology apart from the rest.

GLOBAL REACH Our sulfur technology is currently operating all over the world, including North America, South America, Europe, and Asia. Recently, we performed basic engineering of sulfur recovery units for four refineries in South America and one in the Middle East, each with an MDEA amine tail gas treating unit followed by tail gas incineration. The sulfur recovery technology compliments Foster Wheeler’s heavy oil conversion technologies including delayed coking, and full EPC capabilities. We are also known in the chemicals, petrochemicals and polymers market. From consultancy and small process unit revamps to large integrated grass root complexes, we deliver comprehensive solutions that meet your requirements. We are truly a global engineering and construction contractor, and power equipment supplier adding value with technically advanced services, reliable facilities and equipment. SPONSORED CONTENT

New Burner Showing Double Air Barrel Reach your peak operating and environmental performance with Foster Wheeler’s Sulfur Recovery Technology! Foster Wheeler develops solutions to meet your Sulfur Recovery needs.

CONTACT INFORMATION Foster Wheeler USA 10876 S River Front Parkway, Suite 250 South Jordan, UT 84095 Phone: 801 382 6900 Fax: 801 382 6901 Email: dean_bybee@fwhou.fwc.com www.fwc.com 585 N. Dairy Ashford Houston, Texas 77079 Phone: (713) 929-5500 Fax: (713) 929-5170 Email: info@fwc.com www.fwc.com HYDROCARBON PROCESSING

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SHELL GLOBAL SOLUTIONS

PRESSURISING THE SHELL SULPHUR DEGASSER How a small modification to enable pressurised degassing operations can substantially cut emissions and enhance safety KEES VAN DEN BRAND, Senior Process Engineer, Gas Treating & Sulphur Processes, Shell Global Solutions International BV

Shell Global Solutions has designed a small modification to its sulphur recovery and degassing configuration that can unlock major cuts in sulphur emissions and also enhance safety. The inexpensive adjustment enables compliance with the stringent World Bank standards. The conventional Shell sulphur degassing process, a well-established technology with more than 330 applications worldwide, removes hydrogen sulphide (H2S) and polysulphides (H2Sx ) from the liquid sulphur produced in Claus sulphur recovery units and sends them to the incinerator. However, burning these sulphur compounds increases sulphur dioxide (SO2) emissions, so Shell Global Solutions has investigated the possibility of recycling these gases to the front end of the Claus unit. In the past, Shell has evaluated reducing SO2 emissions by using a compressed recycle towards the main burner with either a compressor or a steam- or air-driven ejector. These options have significant drawbacks. A compressor is expensive, a steam-driven ejector cools the main flame too much, and an air-driven ejector results in poor turndown on the air side of the main burner. The conventional Shell sulphur degasser operates at near atmospheric pressure (see FIG. 1). The sulphur rundown lines from the Claus unit (which also operate at atmospheric pressure) drain by gravity flow into the degasser. Traditionally, this was a concrete pit, but in recent years there has been a design shift away from

Low-pressure steam

the use of pits in favour of vessels. This is a crucial development because having a vessel unlocks the possibility of recycling the vent gas from the degasser to the front end of the Claus unit at a slightly elevated pressure. As the vessel is leak tight, sweep air is not required to prevent uncontrolled leakage. Furthermore, the vessel facilitates more robust safeguarding and tracing solutions for corrosion prevention. This approach also offers the additional flexibility of off-plot installation. The conventional atmospheric configuration requires the degasser to be placed near the sulphur condensers to minimise the pressure drop in the rundown. In contrast, with the new pressurised line-up, the degasser vessel can be installed in another part of the Claus plot, which can result in more efficient use of the available plot space. The new configuration for pressurised sulphur degassing (see FIG. 2) requires a small Roots-type blower to create a pressure of about 0.9 barg in order to recycle the vent gas to the Claus unitâ&#x20AC;&#x2122;s front end, as well as an additional small collecting vessel. Shell carefully evaluated the effect of higher pressures on degassing performance. The tests confirmed that degassing is more effective at a slightly elevated pressure. More oxygen can dissolve in the sulphur, which enhances the decomposition of H2Sx. Therefore, for the same efficiency, the residence time could be decreased or smaller units could be used.

To incinerator Air

Air

Sulphur from Claus plant

Sulphur to storage

Sweep gas

Stripping column with separation baffle

FIG. 1. Simplified flow scheme of the atmospheric Shell sulphur degassing process. S-80

SULFUR 2012 HydrocarbonProcessing.com

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SHELL GLOBAL SOLUTIONS

Vent air to Claus main burner gun

Air

To/from sulphur recovery unit coalescer

Sulphur degassing vessel

Sulphur collecting vessel

Liquid sulphur from sulphur locks

Bubble column

Bubble column Liquid sulphur to storage

Concrete pit

FIG. 2. The pressurised Shell sulphur degassing process.

DELIVERING VALUE World Bank standards specify that no more than 150 mgSO2/Nm3 (about 53 ppmv SO2 ) must leave the incinerator (dry basis and no oxygen). This is typically equivalent to 35 ppmv of SO2 in the actual stack gas. These limits, or even tighter ones, are increasingly being specified on projects all around the world and setting major challenges for operators. The pressurised Shell sulphur degassing process offers a safe, robust and cheap solution for achieving the required reductions in SO2 emissions. The main modifications required for pressurised operations are: • replacing the below-ground concrete pit with a below- or above-ground vessel, if appropriate; • recycling the degasser vent to the front end of the Claus unit; and • installing a simple compressor. These changes require only minor capital expenditure. The advantages include: • reduced SO2 emissions; • enhanced safety; and • plot flexibility. Refineries, gas plants and upstream facilities can all benefit, either in grassroots installations or retrofit situations, and Shell expects the pressurised degasser to become the default configuration in future projects that have highly stringent SO2 emission requirements.

CONTACT INFORMATION www.shell.com/gasprocessing gasprocessing@shell.com

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HOUSTON HOUSTON

2012 WOMEN’S

Gold Sponsor:

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Register online at WGLNetwork.com The Women’s Global Leadership Conference in Energy and Technology (WGLC) is one of the largest women’s events in the industry, and the only one that focuses on discussing key environmental, economic and professional development issues in oil and gas.

Silver Sponsors:

This year, Houston Mayor Annise Parker will deliver a welcome address the morning of Oct. 30 to open day one of the conference. The 2012 WGLC keynote speakers are:

Henrietta H. Fore Chairman and CEO Holsman International

Mark P. Mills CEO Digital Power Group

This year, conference content will focus on the global impact of recent technological advances in exploration and production. Presentations will also cover IOC/NOC relationships, work force trends, shale energy in North America, challenges facing young professionals, global upstream activity, and more.

Bronze/Conference Bag Sponsor:

Additional speakers include: • Melody Meyer, President, Chevron Asia Pacific Operation Company • Cindy Yeilding, VP, Gulf of Mexico Exploration, BP • Alexandra Roberts-Judd, Upstream Projects Public and Government Affairs Manager, ExxonMobil

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Heat Transfer S. AHAMAD and R. VALLAVANATT, Bechtel Corp., Houston, Texas

Identify and control excess air from process heaters Excess O2 can be measured in flue gases, which can be correlated with excess air. FIG. 1 provides a correlation between excess air and flue gas O2 for a typical natural gas. Also, the following equation can be used to calculate the excess air based on flue gas O2 : EA

(1)

where: EA = excess air, % O2 = vol% of flue gas oxygen (dry). Higher excess O2 helps achieve a stable flame in the firebox. At the same time, it reduces the efficiency of the heater. As a general rule, 3% O2 in flue gas is equivalent to 15% excess air. Flue gas quantity increases with a rise in excess air, which

lowers heat and increases the fuel requirement. FIG. 2 provides Excess air vs. O2

40 Excess air, %

Excess air is defined as the amount of air above the stoichio-

metric air requirement that is needed to complete the combustion process. Excess oxygen (O2 ) is the amount of O2 in the incoming air not used during combustion. In an operating plant, the airflow rate can be adjusted at a fixed absorbed-heat duty (constant feed flowrate and inlet/outlet conditions) until an optimum fuel-to-air ratio is achieved. It is important to note that there is a limit on minimum possible excess O2 . Below this level, combustibles can enter the flue gas, which poses a safety hazard. Heater and burner manufacturers establish this minimum limit during the design stage. Operators should also keep a safe margin for upset conditions. A frequently asked question is, “Why do operators often run heaters with higher excess air?” The answer is that additional excess air reduces flame temperature, shortens flame length and decreases tube flame impingement, thereby making it easier for workers to operate the heater without overheating the tube.

92 O2 21 – O2

Flue gas O2, % wet

Process heaters are the largest consumers of energy in most plants. A refinery, on average, burns approximately 2 billion Btu/hr of fuel in fired heaters. The total quantity of fuel burned (heat released) is so high that any improvement will result in significant fuel savings. Although there are many ways to improve heater performance, including better design, operation and maintenance, excess air is the No. 1 contributor to poor heater efficiency and must be addressed. High energy costs and tighter emissions regulations require increased understanding and control of excess air. Any reduction in excess air will raise the efficiency of a heater and reduce total emissions. NOX emissions are of the highest concern in a fired heater, although excess air control will also reduce refinery CO2 emissions and boost heater efficiency. Fuel efficiency in a fired heater is a function of heater design, maintenance and operating parameters. Heaters must be designed for optimum efficiency, and providing an adequate heat-transfer area at the design stage will ensure better efficiency. It is recommended that the flue gas temperature approach (defined as flue gas temperature leaving convection, minus process inlet temperature) be between 50°F and 100°F, depending on heater tube material and the cost of fuel. However, heater efficiency may decline with the degradation of heater components. The degree of degradation is dependent on the quality of the maintenance program implemented at the refinery.

8 7 6 5 4 3 2 1 0

O2—dry vs. wet

4 6 Flue gas O2, % dry

FIG. 1. Correlation between excess air and flue gas O2 for a typical natural gas. Hydrocarbon Processing | OCTOBER 2012 83

Heat Transfer a correlation between flue gas generated during combustion, and excess air. The following equation can be used to calculate approximate flue gas quantity for natural gas: qf = 1 + 0.167 âŤť (100 + EA) or

(2) q f = 17.7 +

15.4 O2 21 â&#x20AC;&#x201C; O2

where: EA = excess air, % O2 = vol% of flue gas oxygen (dry) qf = flue gas quantity in lb/lb of fuel. As a general rule, flue gas quantity is approximately 20 times the fuel quantity at 15% excess air. The net efficiency of a fired heater is equal to the total heat

absorbed, divided by the total heat input. The heat absorbed is equal to the total heat input, minus the total losses. The net thermal heater efficiency can be calculated using the following equation: Total heat input â&#x20AC;&#x201C; (stack + setting) losses 100 Total heat input LHV H a H f Hm â&#x20AC;&#x201C; H s â&#x20AC;&#x201C; H L = 100 LHV H a H f Hm =

(3)

Hf = Sensible heat of fuel, Btu/lb of fuel Hm = Sensible heat of atomizing media, Btu/lb of fuel Hs = Stack heat losses, Btu/lb of fuel HL = Setting loss, Btu/lb of fuel. For all practical purposes, we can assume Ha and Hf to be negligible. Hm is applicable for fuel oil firing. Setting (casing) heat losses are in the range of 1.5%â&#x20AC;&#x201C;2.5%, depending on the capacity, design and size of the heater. Given these assumptions, there are two parameters for the estimation of efficiency: excess air/flue gas O2 and stack temperature. FIG. 3 depicts a graph for the estimation of fired heater efficiency, based on flue gas O2 and stack temperature for a typical natural gas with a setting heat loss of 1.5%. For heat losses higher than 1.5%, additional heat loss should be reduced from the calculated efficiency. For example, consider a heater operating at a stack temperature of 400Â°F, with 4% O2 (dry), and a 1.5% setting loss. Using the graph in FIG. 3, the efficiency can be estimated at 89%. For the same heater with a higher setting loss of 2.5%, the efficiency is 88%. Knowledge of efficiency loss will clarify economic incentives to lower the stack temperature or the percentage of excess O2. As a general rule, every 35Â°F increase in flue gas temperature reduces the heater efficiency by 1%. Natural draft heaters use the draft (buoyancy) effect of hot

flue gases to draw combustion air into the heater. The net draft available is the draft created by the stack effect, minus frictional and velocity losses. The net draft should be sufficient to obtain a negative pressure along the heater flue gas path. It is important to maintain a safe draft level in a fired heater to achieve the best possible efficiency and operation. The

where: Îˇ = Net thermal efficiency, % LHV = Lower heating value of the fuel, Btu/lb of fuel Ha = Sensible heat of air, Btu/lb of fuel Flue gas quantity

Flue gas stack temperature, Â°F 300 400 500 600 700 800 900 1,000

85 Efficiency, %

24 lb of ďŹ&#x201A;ue gas per lb of fuel

Flue gas O2 vs. efficiency

75 20 70 18 65 16

30 Excess air

FIG. 2. Correlation between flue gas generated during combustion, and excess air.

84 OCTOBER 2012â&#x20AC;&#x201A;|â&#x20AC;&#x201A;HydrocarbonProcessing.com

6 8 Flue gas O2 , vol% (dry)

FIG. 3. Estimation of fired heater efficiency.

Heat Transfer target draft of 0.1 inchWC is set at the heater arch. A higher value of draft will result in ingress of “tramp air” into the heater. Tramp air takes heat from the combustion process and exits the stack, reducing heater efficiency. The flue gas sample taken from the stack does not represent the actual volume of O2 available for combustion. It is the sum of unused O2 from the firebox (actual excess O2 ) and O2 from tramp air. A positive draft value will result in the leakage of hot flue gases through openings in the heater. This is a hazardous operation that can overheat the steel structure, refractory and heater supports, and, consequently, shorten heater life. FIG. 4 provides the value of draft generated in the heater for flue gas temperature and ambient air temperature. It should be noted that stack effect decreases with an increase in site altitude. The calculated draft should be amended using the correction factor for site altitude. The following equation can be used to calculate the draft generated in a heater:

Another method of measuring leakage involves heat balance. The flue gas/air heat balance across the APH can be described as follows: (m ⫻ Cp ∆T)flue gas = (m ⫻ Cp ∆T)air

(6)

(4)

For a typical fuel gas at 15% excess air: mflue gas ≈ 1.05 ⫻ mair (7) Cpflue gas ≈ 1.15 ⫻ Cpair ∆Tair ≈ 1.2 ⫻ ∆Tflue gas where: m = Flowrate Cp = Specific heat ∆T = Temperature difference across the APH. Any leakage in the APH will reduce the ratio of ∆Tair to ∆Tflue . For example, for a 10% leakage in the APH, the ratio of temgas perature difference will be around 1.1. FIG. 5 indicates the percentage of air leakage based on the ratio of ∆Tair to ∆Tflue gas for a typical natural gas firing.

where: H = Height, ft Patm = Atmospheric pressure, psia Tamb = Ambient air temperature, °R Tfg = Flue gas temperature, °R. As a general rule, for every 10 ft of firebox height, the draft increases by 0.1 inchWC:

Air leakage through openings. A fired heater is not a 100% sealed unit; there are always openings through which air ingress (tramp air) can move. The volume of tramp air depends on the opening size and the draft at the location of the opening. After the draft at the opening location is estimated, the following equation can be used to estimate the air leakage through an opening:

Draft = 0.52

Patm

1 Tamb

–

1 Tfg

Draft at burner (inchWC) ≈ 0.1 + Hfb ÷ 100

∆P = C ⫻ 0.003 ⫻ ρ ⫻ v 2

(5)

where: Hfb = Firebox height, ft. As an example, for a 50-ft-tall firebox, the draft at burner is 0.6 inchWC (0.1 inchWC at arch, plus 0.5 inchWC as a stack effect). Generally, the stack effect decreases with an increase in site elevation. For example, a taller stack will be needed for a heater operating in Wyoming (altitude ~ 5,000 ft) than for one operating along the Texas Gulf Coast (altitude ~ 0 ft).

Stack effect (draft)

1.2 1.1 1.0 0.9 Draft per 100 ft, inchWC

Ambient air temperature, °F 60 70 80 90 100 110 120 130

0.8 0.7 0.6

1.00

Altitude correction factor

0.95

0.5

0.90 Factor

The typical combustion air preheater (APH) will increase the heater efficiency by approximately 10%. Fuel gas generally contains H2S or sulfur, which convert into SO2 and then into SO3. The APH’s heat-transfer surface is subject to cold-end corrosion caused by condensation of sulfur trioxide (SO3 ), which results in APH leakage. Air preheater leakage is one of the most common APH operating problems, and any such leakage results in a reduction in the overall efficiency of the heater. In APH operation, the flue gas is generally at negative pressure, and the air is at positive pressure. Therefore, leakage occurs from air to the flue gas side. This reduces the quantity of air available for combustion, and it increases the quantity of flue gas leaving the APH. This leakage can be detected by measuring the flue gas O2 content at the APH inlet and outlet. Any leakage will result in higher flue gas O2 at the APH exit, compared to the APH inlet. Generally, the APH is not equipped with a flue gas O2 analyzer at the inlet and the outlet; however, the inclusion of 2-in. connections at the APH’s inlet and outlet will enable operators to measure O2 levels using a portable analyzer.

(8)

0.4

0.85 0.80 0.75 0.70

0.3 0.2 200

700

0 1,000 2,000 3,000 4,000 5,000 6,000 7,000 Altitude, ft

1,200 Flue gas temperature, °F

1,700

2,200

FIG. 4. Value of draft generated in the heater for flue gas and ambient air temperatures. Hydrocarbon Processing | OCTOBER 2012 85

Heat Transfer This equation can be simplified for the leakage calculation purpose based on the following data: â&#x20AC;˘ Molecular weight (MW) of air = 28.96 â&#x20AC;˘ Atmospheric pressure (psia) = 14.7 â&#x20AC;˘ Velocity head (C ) = taken as 1 The simplified equation reads: ql =

â&#x2C6;&#x2020;P

115

(9)

where: Î&#x201D;P = Draft at opening location, inchWC Ď = Density of air at ambient temperature, lb/ft3 v = Velocity of air through opening, ft/s C = Velocity head ql = Air leakage, lb per ft2/s T = Ambient air temperature, Â°R. FIG. 6 provides the quantity of air leakage per ft2 of opening size. This figure is based on an ambient air temperature of 60Â°F. Once the opening size is known, the amount of air leakage can be estimated. The estimated air can be translated into the additional firing rate required. Fuel savings. A commonly asked question in heater discus-

sions is, â&#x20AC;&#x153;How much fuel can be saved if excess air is optimized?â&#x20AC;? The efficiency chart in FIG. 3, which shows operating O2 and target O2, helps calculate the savings. However, there is a drawback. The absorbed heat duty of the fired heater is constant. Any increase in the O2 level will reduce the efficiency, resulting in a higher firing rate. This increase in the firing rate will lead to a rise in stack temperature, which results in another reduction in efficiency. This reduction, in turn, demands a further increase in the firing rate. APH leakage

For example, a 100-MM-Btu/hr fired heater is designed for operating at a stack temperature of 600Â°F, with 84% efficiency at 3% O2. The operating efficiency at 6% O2 is around 80% (and not 82%, as shown in the efficiency chart). The method of efficiency calculation for off-design operating conditions presented in API-560 Appendix G can be used to estimate the stack temperature when excess air is present. This method can be simplified for excess air as follows: TS2

Tf 2

(TS1 T f 1 )

100 EA2 100 EA1

(10)

0.35

180 TS1 T f 1

where: Ts = Flue gas stack temperature, Â°R EA = Excess air, % Tf = Feed inlet temperature, Â°R (Tf1 = Tf2 ) ÎŚ = Excess air correction factor (subscripts 1 and 2 refer to design and operating conditions, respectively). Once the new flue gas stack temperature at excess air is known, then the heater efficiency can be estimated. FIG. 7 shows the estimated fuel savings for a reduction in the O2 level to 3%. This graph is based on a fuel price of $6/MMBtu. The design flue gas temperature lines indicate the baseline stack temperature (i.e., the flue gas stack temperature at 3% O2). CO2 emissions reduction. The volume of CO2 emissions generated in a fired heater is directly proportional to the firing rate. In combustion processes, fuel carbon converts into CO2. Therefore, excess air reduction will lower CO2 emissions. Air leakage

20,000

35 18,000

16,000 Air leakage per ft2 of opening, lb/hr

APH air leakage, %

25 20 15

14,000 12,000 10,000

8,000

6,000

0 0.6

0.7

0.8

0.9 Î&#x201D;T air/Î&#x201D;T ďŹ&#x201A;ue gas

1.0

1.1

FIG. 5. Percentage of air leakage for a typical natural gas firing.

86 OCTOBER 2012â&#x20AC;&#x201A;|â&#x20AC;&#x201A;HydrocarbonProcessing.com

1.2

4,000 0.0

0.2

0.4

Draft, inchWC

0.6

FIG. 6. Quantity of air leakage per ft2 of opening size.

0.8

1.0

Heat Transfer FIG. 8 provides estimated decreases in CO2 emissions through

a reduction in the O2 level to 3%. The basis for this graph is the same as that in FIG. 7.

Recommendations. Heater excess air control starts at the de-

sign stage. Well-designed heaters have low tramp air. There are three stages of excess air control: 1. Design stage 2. Maintenance 3. Control. The following methods can be used to reduce excess air and tramp air in the heater. Design stage. A heater has many potential leak points for air ingress: • Clearance around the bottom coil guide (spigots) • Sight doors and peepholes • Header boxes, manholes and other openings for viewing and access • Modules and duct splice joints • Terminals and crossover tubes • Weld joints on the heater casing • Soot-blower sleeves • The APH. These leak points must be designed for the lowest possible leakage. Suggestions for designing a low-leakage heater include the following: • Seal the clearance space around the bottom tube guides by using a floor sleeve with an end cap, or seal boots • Use sight doors, with safety glass, that are equipped with an interlock cover or flapper • Use a self-closing peephole cover in the heater floor • Ensure that header box panels and other openings are airtight, and use gaskets between the gaps • Seal-weld all splice joints between modules from the

inside, or use high-temperature sealant; also, use closer-bolt spacing (6 in. from center to center) • Seal all terminals and crossover openings with flexible seals • Ensure that all header box drain points are plugged • Ensure that no leakage is occurring through instrument mountings • Limit leakage through the APH during the design stage, and perform an air-leakage test in the shop. Maintenance. Routine maintenance of the heaters is essential, since corrosive agents can be present in flue gases. Deterioration from sulfur oxides occurs mostly on cold sections of the steel casing. Climate conditions can also lead to rusting on exposed surfaces of the heater casing. Suggested inspection and maintenance methods include the following: • Check for heater casing corrosion; if any leaks are discovered, they should be sealed to stop air ingress • Ensure that observation doors (generally located in the bottom section of the radiant box) are closed after technicians inspect the heater flame • Check peepholes, access doors, etc., for proper closing • Check flue gas O2 content in the convection section and on the APH; if there is any increase in O2 content across the flue gas path, it indicates leakage • Use a smoke test during heater shutdown to detect leakage • Use infrared scanning, while the heater is in operation, to pinpoint locations with air leakage; these will have localized, lower heater casing temperatures • To reduce leakage in burners, keep all burners in operation, even during lower operating loads; and close the air register when a burner is taken out of service. CO2 reduction

4,500

Design ﬂue gas temperature, °F

4,000

Fuel savings

2,000,000

3,500 400 600 800

3,000 Annual CO2 reduction, MMlb

Design ﬂue gas temperature, °F 1,500,000 Annual fuel saving, US dollars

400 600 800

Basis: đƫ !/%#*ƫý1% ƫ%*(!0 temperature = 300 °F đƫ 1!(ƫ,.% !ƫœƫĸćĥ 01 đƫ /+. ! ƫ$! 0ƫ 105ƫœƫāĀĀ ƫ 01ĥ$.

1,000,000

Basis: đƫ !/%#*ƫý1% ƫ%*(!0 temperature = 300 °F đƫ 1!(ƫ,.% !ƫœƫĸćĥ 01 đƫ /+. ! ƫ$! 0ƫ 105ƫœƫāĀĀ ƫ 01ĥ$.

2,500 2,000 1,500 1,000

500,000 500 0 0

6 8 Flue gas O2, vol% (dry)

FIG. 7. Estimated fuel savings for a reduction in the O2 level to 3%.

6 8 Flue gas O2, vol% (dry)

FIG. 8. Estimated decreases in CO2 emissions through a reduction in the O2 level to 3%. Hydrocarbon Processing | OCTOBER 2012 87

Heat Transfer Excess air control. Knowing the target flue gas O2 content is the first step in excess air control. Each heater is unique in its design. The O2 level required to achieve ideal combustion may be anywhere from 1%â&#x20AC;&#x201C;4% or higher, depending on the design and operating characteristics of the heater. The following two instruments are necessary to control excess air: â&#x20AC;˘ Flue gas O2 analyzer. This is the most important instrument on the heater. It is recommended to install an O2 analyzer at the radiant section arch. â&#x20AC;˘ Draft gauge. A draft gauge should be installed at the heater arch. The arch is the point of the highest flue gas pressure in the heater. Heater O2 and draft at the radiant arch should be checked and, if necessary, adjusted at least once per shift and whenever there is a change in process load. All operators should be familiar with the heater controls. Often, heaters with air registers and stack dampers become jammed simply because they are not used. FIG. 9 provides tactics for controlling excess air in a natural draft heater. For controlling excess air in other types of heaters, TABLE 1 and TABLE 2 can be used alongside FIG. 7. Automatic control. The basics of automatic heater control are

similar to those described in the manual control method. In automatic control, reliable instrumentation is key. One reason for the small number of heaters with automatic control is a lack of confidence in reliable instrumentation. Also, artificial intelligence can be built into the control system to account for all operating cases. A remote manual control for both O2 and draft is best suited for the natural draft heater. A fully automated control can be safely implemented on balanced draft heaters. Optimum heater performance can be achieved by controlling O2 and combustibles in flue gas using an O2/CO analyzer and automatic dampers.

Case Study 1. In this case study, a vertical, cylindrical, natural draft heater with an absorbed heat duty of 100 MMBtu/hr included the following design parameters: â&#x20AC;˘ Flue gas stack temperature: 600Â°F â&#x20AC;˘ Fluid inlet temperature: 300Â°F â&#x20AC;˘ Design excess air: 15% â&#x20AC;˘ Design heat loss: 2% â&#x20AC;˘ Firebox height: 55 ft â&#x20AC;˘ Number of radiant tubes: 64 â&#x20AC;˘ Flue gas O2 content at operating conditions: 6 vol% (dry) â&#x20AC;˘ Efficiency of heater at design conditions: 84.2% at 1.5% heat loss (see FIG. 3) â&#x20AC;˘ Efficiency at 2% heat loss: 83.7% â&#x20AC;˘ Firing rate: 100/0.837 = 119.5 MMBtu/hr â&#x20AC;˘ Fuel gas flowrate: 5,770 lb/hr (using a fuel gas LHV of 20,700 Btu/lb) â&#x20AC;˘ Flue gas flowrate: 5,770 âŤť [1 + 0.167 âŤť (100 + 15)] = 116,582 lb/hr â&#x20AC;˘ Air flowrate (flue gas flowrate â&#x2C6;&#x2019; fuel flowrate): 116,582 â&#x2C6;&#x2019; 5,770 = 110,812 lb/hr. The heater operated under the following conditions: â&#x20AC;˘ Flue gas O2: 6 vol% (dry) â&#x20AC;˘ Excess air: (92 âŤť 6)/(21 â&#x20AC;&#x201C; 6) = 36.8% (also see FIG. 1) â&#x20AC;˘ Efficiency at higher excess air, design stack temperature of 600Â°F and 2% heat loss = 81.3%; however, the revised efficiency must be calculated based on operating stack temperature, and if this measurement is not available, then the revised stack temperature can be estimated as follows:

TS2 n

TABLE 1 . Methods for controlling excess air O2 level

Natural/induced draft

Forced/balanced draft

High

Close burner air register

Close fan/duct air damper

Low

Open burner air register

Open fan/duct air damper

TABLE 2. Methods for controlling draft Draft level

Natural/forced draft

Induced/balanced draft

High

Close stack damper

Close induced-draft fan damper

Low

Open stack damper

Open induced-draft fan damper

TS2

Tf 2

(TS1 T f 1 ) 0.35

^600180300 `

100 36.8 100 15 300 1.156

0.836

(11)

0.836

1.156 (600 300)

647qF

â&#x20AC;˘ Efficiency at revised stack temperature and operating flue gas excess air: 80.5% (from FIG. 3) â&#x20AC;˘ Corrected efficiency for 2% heat loss: 80% â&#x20AC;˘ Operating firing rate: 100 Ăˇ 0.8 = 125 MMBtu/hr â&#x20AC;˘ Fuel savings potential: (125 â&#x20AC;&#x201C; 119 .5) âŤť 24 âŤť 365 = 48,180 MMBtu/yr; at a fuel price of $6/MMBtu/hr, the annual fuel savings potential = $289,000 (this can also be estimated using FIG. 7). Case Study 2. In this case study, the heater experienced leak-

Close stack damper High

Start

Draft

Low

Close air registers

High O2

OK Open air registers

Low

Open stack damper

FIG. 9. Tactics for controlling excess air in a natural draft heater.

88 OCTOBER 2012â&#x20AC;&#x201A;|â&#x20AC;&#x201A;HydrocarbonProcessing.com

Finish

age through the bottom guide and an open peephole. The coil guide and sleeve size were 2-in. Nominal Pipe Size (NPS) Schedule 80 and 3-in. NPS Schedule 80, respectively. The heater was not provided with a cap on the sleeve. The open area between the sleeve and the guide included an inside cross-section of 3-in. NPS sleeve and an outside cross-section of 2-in. NPS guide (equal to 2.96 in.2). Other design parameters included: â&#x20AC;˘ Number of guides: 32 (one guide for two radiant tubes) â&#x20AC;˘ Total opening area at guides: 32 âŤť 2.96 Ăˇ 144 = 0.66 ft2 â&#x20AC;˘ Size of radiant section observation door: 5 in. âŤť 9 in. â&#x20AC;˘ Opening area of observation door: 45 in.2 (0.31 ft2) â&#x20AC;˘ Total open area: 0.97 ft2

Heat Transfer â&#x20AC;˘ Firebox height: 55 ft â&#x20AC;˘ Draft at opening (draft at heater floor): 0.1 + 55 Ăˇ 100 = 0.65 inchWC â&#x20AC;˘ Ambient air temperature: 60Â°F â&#x20AC;˘ Air leakage per ft2 of opening: 115 0.65 = 4.07 lb/s/ft 2 (460 + 60) â&#x20AC;˘ Total air leakage: 4.07 âŤť 0.97 âŤť 3,600 = 14,212 lb/hr (this can also be estimated using FIG. 6) â&#x20AC;˘ Excess air for the burner: 15% â&#x20AC;˘ Actual excess air, including leakage, can be calculated as follows: (Air leakage + air required) Air required (14,212 110,182) 110,182

(100 EA) 100

(100 15) 100

29.8%

Based on revised excess air, a fuel savings can be calculated as in Case Study 1: â&#x20AC;˘ Calculated efficiency: 81.1% â&#x20AC;˘ Loss in efficiency due to air leakage: 83.7 â&#x2C6;&#x2019; 81.1 = 2.5% â&#x20AC;˘ Annual fuel savings potential = $201,320. Recommendations. Controlling excess air has many benefits, and opportunities exist to save fuel regardless of whether

the heater is old or new. Day-to-day monitoring significantly improves heater operation. The first visible benefit of excess air reduction is a decrease in fuel consumption. Reduction of emissions, including CO2, is another benefit. The heater must be provided with at least two instruments (an O2 analyzer and a draft gauge at the arch), both of which are important for energy improvement. Additionally, proper training should be provided for heater operators, and a simple and straightforward heater tuning program must be implemented. It is unlikely that an operator will voluntarily adjust the heater unless a plan is in place to do so. SULTAN AHAMAD is a fired-heater equipment engineer at Bechtel Corp. in Houston, Texas. He has more than 14 years of experience in the design, engineering and troubleshooting of fired heaters and combustion systems for the refining and petrochemical industries. He graduated from the Indian Institute of Technology in Roorkee, India, with a degree in chemical engineering. He worked for eight years at Engineers India Ltd. in New Delhi, India, and for five years at Furnace Improvements in Sugar Land, Texas. RIMON VALLAVANATT is the senior principal engineer at Bechtel Corp. in Houston, Texas. He has more than 37 years of experience in the design, engineering and troubleshooting of fired heaters, thermal oxidizers, boilers and flares. He graduated from the University of Kerala in India with a degree in mechanical engineering. He also received a degree in industrial engineering from St. Maryâ&#x20AC;&#x2122;s University in San Antonio, Texas. Mr. Vallavanatt is a registered professional engineer in the state of Texas, and he has served on the American Petroleum Instituteâ&#x20AC;&#x2122;s subcommittee on heat transfer equipment for the past 27 years.

INDUSTRIAL NITROGEN

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Hydrocarbon Processingâ&#x20AC;&#x201A;|â&#x20AC;&#x201A;OCTOBER 2012 89

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BECAUSE IT IS ALL GLASS: s )T WILL NOT ABSORB mAMMABLE LIQUIDS s )T WILL NOT BURN s 7ILL NOT PROMOTE mAME SPREAD s 7ILL NOT PRODUCE TOXIC SMOKE OR OFF GASSES

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foamglas.com/industry

Rotating Equipment K. BIHLER, Bihlertech, Inc. Chicago, Illinois; D. DOMINIAK, Automate the World, Inc., Chicago, Illinois; B. KEITH, Toshiba International Corp., Houston, Texas; and J. JOHNSON, Hydro Inc., Chicago, Illinois

Apply new pump-drive software to test performance Process control schemes involving fluid-flow machinery with traditional adjustable speed drives (ASDs) or modules encounter limitations such as control flexibility and accuracy. In this context, the traditional approach implies that proportional integral derivative (PID) methods can be used and that these methods can directly control frequency. However, applying PID and directly controlling frequency on centrifugal pump drives restricts flowrates and pressures along a nonlinear pump-performance curve. The widely known pumpperformance curve shapes are determined by a particular machine’s impeller design or fabrication details. With respect to controlling flowrate and/or pump discharge pressure, ASD modules using traditional PID methods suffer from trying to apply a linear equation to solve a nonlinear problem. In contrast, a newly developed pump performance algorithm (PPA) can indirectly solve for power in Eq. 1 of the Pump Affinity Laws. As shown in FIG. 1, a PPA was developed for modern centrifugal pump applications and test-stand duties where superior control is needed. The superior control is achieved when an algorithm “linearizes” a pump curve; it offers expanded new ASD-related control options and improved energy efficiency as compared to conventional PID methods. Extensive testing corroborated the developments, and testing results will be described in this article.

However, when the target speed is reached using this method, an exact answer is not available and the desired setpoint may have been missed. The correct answer lies on an unknown and nonlinear pump performance curve. PID software must now guess the speed and, hopefully, decrease the error. This process continues until a reasonable closeness to (or deviation from) setpoint is reached. If any event or process variable causes the system curve to change, the whole control process must be reattempted. The process of searching for the correct speed will consume driver energy, even on a single machine. Multiple machines. On a fluid-flow system or fluid-processing unit with multiple machines operating either in series or in parallel, a second and more serious problem—balancing the load—presents itself. Even if the machines in a multi-machine flow loop or fluid-moving unit are built to the same specifications, there will be differences in motor performance, impeller clearances, internal wear, surface roughness, and perhaps other parameters. Some or all of these conditions can affect the relationship (or constancy and accuracy) of pump speed and its associated head vs. flow performance. Using conventional frequency control, multiple fluid-flow devices (pumps) are operated at the same speed when running together. For the reasons listed above, this will inevitably pro-

Pump laws and equations. Consider the Pump Affinity

Laws (Eq. 1), where the diameter of the impeller is held constant and: Q = Flow P = Power H = Pressure N = Impeller speed Flow Q1 Q2

N1 N2

Pressure H1 H2

N1 N2

Power 2

P1 P2

N1 N2

(1)

Note: The basic principle behind PPA modules and software used in these devices solve for power, P and indirectly solve for speed, N.

COMPARING OLD VS. NEW CONTROL METHODS Common PID software used in an ASD provides an estimated speed to achieve a desired or targeted head vs. flow setpoint.

FIG. 1. Modern pump test stand with ASD drive modules (cabinets) on the right side of the picture. Photo courtesy of Hydro Inc., Chicago, Illinois, and Toshiba International Corp., Houston, Texas. Hydrocarbon Processing | OCTOBER 2012 91

Rotating Equipment the testing and two pumps were operated in parallel and series configurations. The various testing sequences involved both standard ASD control and also the new PPA. During a straightforward test with a single pump, many of the different capabilities of PPA software were demonstrated. This software maintained constant system pressure as pumps were added or shut down to modulate flow, or, as was done in FIG. 2, by modifying the flow through a single pump by partially closing a discharge valve and purposefully altering the pump speed. Likewise and by holding a particular PPA number, the specifically developed software maintained constant power usage. Power usage corresponds to the PPA number by altering the frequency in response to externally imposed flow changes. Each pump is tuned during The advanced PPA software balanced loads and eliminated the process of searching for a setpoint. commissioning to produce a specific The PPA starts with a fixed PPA number. If the outresult, such as a flowrate or pressure. put electrical current is less than the PPA number, then the frequency is increased. Otherwise, if the Each motor/pump combination is unique output current is greater than the PPA number, then and requires slightly different frequencies the frequency is decreased. The frequency (in Hz) is directly proportional to driver rpm (N in Eq. 1) and power requirements. and is adjusted to meet the driver’s electric current (amperage, or amp) draw. On centrifugal process pumps, amp draw is linear to both pressure and flow, and nonlinear to N. This then differs from the PPA algorithm, TEST OVERVIEW AND PPA DEMONSTRATION which controls power directly to meet pressure, flow or a choAs part of the development work, the deficiencies of simple sen pressure vs. flow intersection. speed control algorithms were addressed; the difference between traditional ASDs and new PPAs had to be quantified. Extensive testing at a state-of-the-art facility (FIG. 1) were SETTING UP PPA scheduled and conducted.1 The results were analyzed and The PPA number is a percentage of the total amperage formally reported. It was ascertained that the testing facility available from a specific ASD. When a motor/pump is first set was designed in compliance with the Hydraulic Institute and up in the PPA wizard, the user is instructed to enter the motor API-610 standards. Instrumentation included a modern datafull-load amps into the drive. The ASD temporarily sets the collection system with electronic flowmeters, torque meters “PPA Max” variable to the entered motor amps divided by the and remote valve controls. Modern PPA drives were used in ASD available output amps. The operator can then adjust the “PPA Max” number to achieve the maximum desired setpoint 1,800 for flow, pressure, etc. Once this is done, the operator will set Total, ﬂow the “PPA Min” variable to find the lower limit, or motor stall 1,600 System, psi point. At this time, the ASD has stored the limits within which 1,400 the pump is allowed to operate. On multiple pump systems, this is repeated so each machine has the same operational set1,200 points. The PPA numbers, which also affect the corresponding 1,000 amperage draw and frequencies at any given operating point, will differ among pumps. Mechanical differences in piping, 800 impeller wear, or trim and motor efficiency are responsible for 600 the differences in PPA numbers. Flow, gpm

duce different flows and pressures from each machine. Result: While in parallel operation, one pump will (usually) produce most of the output and require a correspondingly higher electric current flow or amperage. Meanwhile, the other pump or pumps will demand a lesser, but still high-power draw while adding very little to the total fluid-flow output. Using conventional control approaches, it will not be possible to balance a load using frequency directly. Again, this inability exists because even the smallest of physical differences will cause a shift in proportional power demand to the fluid device with the highest pressure-to-flow ratio.

400

Virtual pump curves. Next, a virtual linear pump curve is

0 -200 -400

884 1767 2650 3533 4416 5299 6182 7065 7948 8831 9714 10597 11480 12363 13246 14129 15012 15895 16778 17661 18544 19427 20310 21193 22076 22959 23842 24725 25608 26491 27374 28257 29140

Pressure, psi

200

FIG. 2. Testing verified that PPA software is capable of maintaining constant pressure even if wide flow variations occur. Source: Hydro, Inc., Chicago, Illinois.

92 OCTOBER 2012 | HydrocarbonProcessing.com

created in software; it describes a power percentage that will be compared to the other pumps. The PPA makes it possible to compare power draws and make speed adjustments until all power draws match each other. If controlled from a common external analog signal, a PPA Min of 0% will cause all the pumps to operate at the PPA Min setpoint. The same is true at 100%, where all pumps will be running at their individual PPA Max setpoints. What is important is that, for each change in setpoint, the resulting flow or pressure will be linear, whereas the frequencies and amp

Rotating Equipment

120

Pressure, psi Hz Hz

Pressure, psi

100 80 60 40 20 0

10:56:54 10:57:19 10:57:43 10:58:07 10:58:31 10:58:56 10:59:20 10:59:44 10:00:08 10:00:33 10:00:57 10:01:21 10:01:45 10:02:09 10:02:34 10:02:58 10:03:22 10:03:46 10:04:11 10:04:35 10:04:59 10:05:23 10:05:48

-20

140 120 100 80

kW kW Total

60 40 20 0 -20 10:56:54 10:57:21 10:57:47 10:58:13 10:58:39 10:59:06 10:59:32 10:59:58 11:00:24 11:00:51 11:01:17 11:01:43 11:02:09 11:02:36 11:03:02 11:03:28 11:03:55 11:04:21 11:04:47 11:05:13 11:05:40 11:06:06

Tuning. Each pump is tuned during commissioning to produce a specific result, such as a flowrate or pressure at both PPA Min and PPA Max. Each motor/pump combination is unique and requires slightly different frequencies and power requirements at both end points and at every point in between. As a percentage, they are all the same. Since all PPA-equipped ASDs share the same common reference signal, they can each independently calculate the power needed to match the flowpressure intersect of all of the pumps online. In a series operation, this effect is even more pronounced. If series pumps are run at the same speed, then they have different fluid velocities.

140

Hertz

Calculating pump speed. Another relevant characteristic of PPA is that, for each ASD actually operating, the software can calculate at what speed to run to hold the PPA number independently. This would be impossible for separate drives running speed-control PID because each drive would be fighting the other while hunting errors until the whole system became unbalanced. Because PPA software provides a virtual and linear line of performance, each ASD can simply select an individual point that represents exactly the same flow-pressure intersect as every other ASD. This means that the total control system is greatly simplified. Therefore, an external PLC or DCS system needs only to tell each PPA-equipped ASD when to run and what setpoint to hold, and to provide the feedback reference signal. The important duty of calculating the specific power requirements to maintain a setpoint is distributed across all of the individual machines that are online. In this scenario, each PPA-equipped ASD that is online, whether the driver is operating or not, is continuously calculating what power it needs to maintain the desired setpoint based on its initial stored setup values. When called, the next drive will ramp up and recalculate the feedback to find its point on the performance line. At the same time, the lead pump will be backing down because the additional flow will be changing its reference signal. Eventually, all pumps will reach the needed frequency to meet the virtual pump performance setpoint, while the PPA software runs in the background to find the required power. This process occurs very quickly because each estimate finds a known performance point. It generally happens faster than the mechanical force variations can cause any pressure or flow changes. This leads to the most important function of PPA software: balancing the load between multiple pumps.

This is analogous to a traffic jam on an expressway where the cars are stopping and starting. Using PPA, the “cars” are all moving at exactly the same velocity. Without PPA, differences in velocities between pumps can result in overpressure waves that amplify and resonate. This causes mechanical stress, which, in turn, consumes more power. With PPA, the frequency can range to satisfy changes in the system curve. An example of this may be found in a common lift station connected to a force main that is shared by other pumping stations. In this case, a transducer would provide feedback to keep the wet well at a constant level. As the flowrate changes, PPA would increase and decrease power to achieve this, just as a standard ASD would range the frequency. In contrast, the different PPA machines solve for P of the Power Affinity Laws simultaneously. At a given power setting on a pump with an increasing power curve, PPA increases the frequency to maintain the specified power when the force main pressure increases as other stations come online. When running in PPA “process hold mode,” the PPA software will do it independently, and, each time the feedback changes the PPA output number. The result is that, unlike with a simple speedcontrol system, the PPA software corrects the frequency based on power before any change comes from the feedback loop. This has a two-fold advantage: more accurate process control and, consequently, less power used by the pump.

Power, kW

draws will be nonlinear. Whenever the input frequency to a single machine is changed in a standard ASD, the output is nonlinear, and energy is wasted searching to find the speed needed to satisfy the control loop. Because PPA produces a virtual linear performance curve, even the first calculation from an external PID controller will likely be the correct and most precise calculation. The PPA software apportions power to the motor to produce the desired pressure or flow. No time or energy is wasted accelerating and decelerating past the needed frequency. When the internal PPA process hold loop is used with an external feedback signal from a pressure transducer or flowmeter, the control loop process is greatly enhanced because there is no communication lag, and the variables can be tested and satisfied at CPU clock speed.

FIG. 3. Two pumps (on common headers), using PID. Their respective speeds lag. Also, maintaining the setpoint pressure (95 psi) is difficult. Hydrocarbon Processing | OCTOBER 2012 93

Rotating Equipment MORE ABOUT THE TESTS Among the tests conducted was Test A, also known as the PID lead-pump test. It involved settings and findings summarized as: • 95 psi setpoint (blue pump) • Speed follower lag pump (red pump) • Average kW usage from 10:59:32 to 11:05:49 = 70.88. As shown in FIG. 3, this test shows the traditional method of using two pumps on a common header. A PID algorithm with feedback is used on the first pump. (Note: PID cannot be used on more than one pump in a system.) The second pump is set to follow the speed reference from the lead machine. The response lag from FIG. 3 is attributable to the acceleration time. With traditional PID, significant energy is wasted in the unbalanced systems. Although operating in parallel, one pump will take over and produce the majority of the flow when all pumps are running at the same speed. This occurs whether on an ASD or across the line. All of the other pumps will draw a considerable amount of power, but will contribute little to the total flow. By contrast, in a PPA system, each machine will precisely produce its portion of the flow and use the same amount of energy. With PPA, all of the pumps in the system become one “virtual” machine.

Test B involved settings and findings labeled Independent PPA operation: • 95 psi setpoint • Average kW usage from 11:57:30 to 12:11:19 = 65.93. As shown in FIG. 4, this test, i.e., Test B, uses the same setpoints as Test A, but places both drives in independent PPA process hold modes. Of note is the smoothness in pressure and flow, and that the drives do not get out of synchronism and fight each other, as shown in the earlier Test A. One additional point of interest is that, in Test A, the PID loop was tuned by a professional application engineer. By contrast, in Test B, the drives, as shown in FIG. 4, were simply operated through the PPA setup wizard, and no tuning was done. Instead, all the values were the drive defaults. From FIG. 4 and comparing it to Test A (FIG. 3), it is clearly demonstrated how PPAs can save energy. Precise control of the process and balancing of the load reduces energy demand that is otherwise converted into heat and mechanical stress. 140

120

Pressure, psi Hz Hz

140 120

Pressure, psi

100 Pressure, psi

100

Pressure, psi Hz Hz

80 60

20 Booster start Hertz

Hertz

20 0 -20

09:20.7 10:15.8 11:10.9 12:05.9 13:01.0 13:56.1 14:51.2 15:46.3 16:41.3 17:36.4 18:31.5 19:26.6 20:21.7 21:16.7 22:11.8 23:06.9 24:02.0 24:57.1 25:54.2 26:49.3 27:44.3 28:39.4

11:29:51 11:32:01 11:34:10 11:36:20 11:38:29 11:40:39 11:42:49 11:44:58 11:47:08 11:49:17 11:51:27 11:53:37 11:55:46 11:57:56 12:00:05 12:02:15 12:04:25 12:06:34 12:08:44 12:10:53 12:13:13 12:15:13

-20

140

80 60 40 20

11:29:51 11:31:52 11:33:53 11:35:54 11:37:55 11:39:56 11:41:57 11:43:58 11:45:29 11:48:00 11:50:01 11:52:02 11:54:03 11:56:03 11:58:04 12:00:05 12:02:06 12:04:07 12:06:08 12:08:09 12:10:10 12:12:11

0 -20

FIG. 4. Test B, same setpoints as Test A but using PPA. Note smoothness.

94 OCTOBER 2012 | HydrocarbonProcessing.com

120 100 80

kW kW Total

60 40 20 0 -20

09:20.7 10:15.8 11:10.9 12:05.9 13:01.0 13:56.1 14:51.2 15:46.3 16:41.3 17:36.4 18:31.5 19:26.6 20:21.7 21:16.7 22:11.8 23:06.9 24:02.0 24:57.1 25:54.2 26:49.3 27:44.3 28:39.4

Power, kW

100

kW kW Total

Power, kW

120

FIG. 5. PPA on one of two pumps operating in series. Header pressure is set at 95 psi.

Rotating Equipment

1,000

400

0 553 1657 2761 3865 4969 6073 7177 8281 9385 10489 11593 12697 13801 14905 16009 17113 18217 19321 20425 21529 22633 23737 24841 25945 27049 28153 29257 30361 31465

Power, kW

200

-200

Pressure, psi

80 60 40

FIG. 6. Two pumps in PPA direct mode reacting to changing discharge header pressures imposed by manipulating a traditional control valve.

06:36.8 06:44.9 06:52.9 07:01.0 07:09.0 07:17.0 07:25.1 07:33.1 07:45.2 07:57.2 08:21.4 08:49.5 09:05.6 10:30.1 10:58.2 09:40.6 10:33.1 11:03.4 11:19.6 11:23.6

Power, kW

FIG. 7. Pump startup without restriction on flow is made possible by PPA. 1,400 1,200 1,000 800 600 400 200 0 -200

Pressure, psi Flowrate

-400

120 100 80 60 40 20 0 -20

kW kW Total

28:19.2 28:48.3 29:17.4 29:46.5 30:15.6 30:44.7 31:13.8 31:42.9 32:12.0 32:41.1 33:10.2 33:39.3 34:08.4 34:37.5 35:06.6 35:35.7 36:04.8 36:33.9 37:03.0 37:32.1 38:01.2 38:30.3

600

Power Power Flow Flow

100

Power, kW

Flowrate, gpm

800

Pressure, psi Hz Hz

120

28:19.2 28:48.3 29:17.4 29:46.5 30:15.6 30:44.7 31:13.8 31:42.9 32:12.0 32:41.1 33:10.2 33:39.3 34:08.4 34:37.5 35:06.6 35:35.7 36:04.8 36:33.9 37:03.0 37:32.1 38:01.2 38:30.3

important pump control issues addressed by PPA software are illustrated in FIG. 7. First, PPA will limit pump overload that normally occurs when starting a pump on an empty pipeline. In this case, the pump having no restriction on flow would electrically overload. The PPA algorithm will, however, limit the output frequency to limit the output amperage. Also, precise process control eliminates wasted energy. This is the key to en-

140

Flowrate, gpm

Direct-mode load balance with an empty pipeline. Two

ergy savings. PPA cannot change pump efficiency or process piping geometry, but it can limit wasted power due to mistakes in the process control introduced by traditional speed control. Second, this test demonstrated that the PPA software is also capable of achieving identical power consumption by two pumps operating in parallel and feeding into a zero back-pressure downstream environment—an empty pipeline.

Pressure, psi

PPA with across the line booster vs. load share. The blue pump drive in FIG. 5 was configured for PPA operation to control pressure at a setpoint of 95 psi. The pressure setpoint was chosen because it was obtainable by the 10-hp pump and motor, but it would cause the pump to operate at an undesirable efficiency level. The red pump drive was configured to emulate a soft-start device—accelerating to 60 Hz, using a 7.5 second acceleration time. At the top chart, the green line shows the pressure on the common header, while the blue and red lines show the output frequency on the respective drives. The lower chart indicates kW consumed for each drive, as well as total kW. In this example, the total power consumption (with both pumps running) is 118 kW. FIG. 5 confirms, in the case of the pumps operating in series and running one at a fixed-line frequency is 36% less efficient than an ASD with PPA for each pump. Operating one pump at a fixed-line frequency and one with an ASD with PID software would be 31% less efficient than an ASD with PPA for each pump. The initial incremental cost of installing an ASD with PPA for each pump would be quickly returned. Inefficiencies will manifest themselves in the form of heat loss, vibration energy and mechanical stress. PPA software will, as in this test case, be about 5% more energy efficient than a traditional speed-control PID. One ASD with PPA for each pump represents the lowest cost of ownership over the service life of the facility. The test results, as shown in FIG. 6, demonstrate both pumps in PPA direct mode reacting to changes imposed by increasing or decreasing the system pressure. The power does not change, and the flows are balanced. In operations similar to this but without PPA, one pump would use most of the power and produce the majority of the flow; the second pump would use somewhat less electric power but contribute very little to the total flow.

FIG. 8. In the direct-mode PPA series both the red-lined and blue-lined pumps consume approximately the same amount of energy. Note: The wide variations in flow and system pressure allowed in this test. Hydrocarbon Processing | OCTOBER 2012 95

Rotating Equipment Series direct mode. As illustrated in FIG. 8, the PPA series direct-mode test shows that, with PPA and system pressure increasing, the power remains balanced. Pressure is represented as the jagged bottom green line in FIG. 8, while power draws for each of the two pumps remain close to identical (red and blue lines). The total power consumed is shown as the black line on the kW vs. time graph.

MORE EFFICIENT PUMP OPERATIONS Over time, pump manufacturers have improved the fluid mechanics in pump designs via increasingly more sophisticated computer CAD/CAM modeling and automated machining capabilities. This is the first advancement step. For the second advancement step, standard ASDs represent notable advantages or innovations. In particular, ASDs achieve better energy efficiency because plain kW usage drops off along with the voltage/frequency ratio. PPAs represent another quantum step forward; it is potentially the third major development benefiting industry. The PPA algorithm enables solving indirectly for the power draw of each individual machine, whether the machine(s) operate alone or in series. Solving indirectly for power in the Pump Affinity Laws or in parallel has been missing from all conventional or traditional ASD methods. The new state-of-the-art PPA software addresses this issue and solves the problem. Enabled by the PPA algorithm, a modern pump drive can quickly reach the desired setpoint. Among the many important

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gains are enhanced energy efficiency and more rapid configuration of complex systems. In addition, operator training is simplified. There is reduced damage potential for fluid machines. Process control via traditional ASDs with centrifugal devices is encumbered by attempting to use a linear equation to solve a nonlinear problem. By contrast, PPA software corrects the power needed before any change is announced from the feedback loop. This rapid before-the-fact correction yields a two-fold advantage: more accurate process control and, consequently, less power consumed by the pump. 1

NOTES Tests conducted at the Hydro, Inc., facilities in Chicago, Illinois. Extensive testing conducted at Hydro accurately measured every relevant parameter and then assisted in defining the important differences between traditional approaches and the PPA algorithms. All involved parties reached the conclusion that PPA software offers unprecedented savings to owner-operators wishing to pursue optimized and reliability-focused pump control.

ASD CAD DCS P PID PLC PPA

NOMENCLATURE Adjustable speed drive Computer aided design Distributed control system Power, as represented by the Pump Affinity Laws Proportional integral derivative Programmable logic controller Pump performance algorithm (a newly developed linear pump performance algorithm)

ACKNOWLEDGMENT This article was edited and refocused by Hydrocarbon Processing’s Equipment and Reliability Editor, Heinz Bloch, P.E. His latest books, Pump Wisdom and the widely read co-authored Pump User’s Handbook (3rd Ed., 2010), are among his 18 comprehensive reliability textbooks. KURT BIHLER has over 30 years of experience in industrial controls. His past work experiences includes programming in the US and Japan for SORD Computer Corp. Since 1993, he has been president of Bihlertech, Inc., a company that specializes in building pump station controls. Mr. Bihler is presently a consultant to Toshiba, and is one of the co-inventors of an advanced PPA algorithm. DANA DOMINIAK holds a PhD in computer science from the Illinois Institute of Technology. With programming experience in many languages, including C and C++, she specializes in advanced automation programming. In addition, she has extensive experience programming graphical interfaces of energy analysis systems for institutions such as Argonne National Laboratory and the International Atomic Energy Agency. Dr. Dominiak serves as an adjunct professor at Lewis University, where she teaches courses in computer graphics and C/C++ Programming. BRIAN KEITH has over 25 years of experience in industrial automation and controls. Starting in the pharmaceutical industry, he has studied under Dr. Melvin First at the Harvard School of Public Health for specialized biological, chemical and radiological containment systems. In 2007, he completed the Georgia Tech Executive Management Training Program. Over the past 15 years, he has focused mainly on the development and marketing of adjustable speed drives for both industrial and commercial markets. Mr. Keith was instrumental in the development and implementation of PPA Technology into the Toshiba ADS family. JEFF JOHNSON has 36 years of experience in the pump service industry covering numerous markets. Having worked with major OEMs around the world including, Sulzer and Flowserve, he joined Hydro in 2009 and was appointed vice president of Hydro’s petroleum and pipeline division. Supporting pump users nationwide, he was instrumental in the design and construction of Hydro’s 5000 HP Test Lab.

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Company

Page

RS#

Website

(77)

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(67)

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Flexitallic LP .............................................................. 5

(65)

Foster Wheeler .............................................S-78–S-79

(96)

FourQuest Energy..................................................... 24

(53)

Gulf Publishing Company Construction Boxscore...........................................S-72 Events—EMGC .............................................S-68, S-73 Events—WGLC ..................................................... S-82 HPI Market Data 2013 ............................................S-71 HPI Marketplace ............................................... 98–99 Hydro, Inc................................................................ 28

www.info.hotims.com/41433-65

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Axens .....................................................................104 www.info.hotims.com/41433-53

(154)

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Bryan Research & Engineering ..................................40

(71)

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Burckhardt Compression AG .......................................13

(79)

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Cameron ................................................................. 50

(55)

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(153)

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(93) (83)

CB&I .............................................................S-74–S-75

(97)

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(157)

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(166)

LAR Process Analysers .............................................. 49

(63)

Linde Process Plants ................................................ 103

(68)

Michell Instruments Inc. ........................................... 54

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Emerson Process Mgmt (DeltaV) ................................. 2 www.info.hotims.com/41433-63

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(61) (62)

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Pittsburgh Corning Corporation................................. 90

(91)

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(160) (161)

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Servomex Ltd............................................................21

(156)

www.info.hotims.com/41433-156

(72) (162) (69)

Shell Global Solutions ................................... S-80–S-81 Sherwin Williams ......................................................32 Spraying Systems Co .................................................. 8

www.info.hotims.com/41433-151

(163)

www.info.hotims.com/41433-163

(85)

(66)

www.info.hotims.com/41433-66

(165)

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(94)

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RS#

www.info.hotims.com/41433-160

www.info.hotims.com/41433-162

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RC Systems .............................................................. 26

www.info.hotims.com/41433-72

(86)

Colfax Americas ........................................................14

(159)

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HyTorc ......................................................................37

Page

www.info.hotims.com/41433-61

www.info.hotims.com/41433-83

Ametek Process Instruments ..................................... 60

Cudd Energy Services ............................................... 89

Flexim Americas Corp. .............................................. 20

Company Website

www.info.hotims.com/41433-93

AMACS ......................................................................25

Chemstations Inc.......................................................22

RS#

www.info.hotims.com/41433-155

Air Products & Chemicals Inc. .....................................27

Cashco Inc. ...............................................................12

Page

Website

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BIC Alliance...............................................................17

Company

(152)

www.info.hotims.com/41433-152

Winsted Corporation .................................................23

(158)

www.info.hotims.com/41433-158

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(89)

www.info.hotims.com/41433-89

www.info.hotims.com/41433-85

(164)

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(92)

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Water Management

LORAINE A. HUCHLER, CONTRIBUTING EDITOR Huchler@martechsystems.com

Consider software tools for water reuse projects The hydrocarbon processing industry needs reliable and economic sources of water for present and future operations. Ensuring a sustainable water supply requires a focused effort to evaluate the quality and quantity of alternative water sources, reuse of individual water streams or the combined outfall stream, and/or changing regulatory requirements. Modeling. The most common method-

ologies to analyze water circuits are pinch technology and mass balance/solution modeling. Application of classic pinch technology for water systems, as shown in FIG. 1, evaluates only the hydraulic demands for water. Better option. A better solution is to construct a sophisticated computerized model of the facility’s water systems that incorporates hydraulic information, along with ionic equilibria of soluble contaminants. Like pinch technology, this approach requires an accurate water balance. FIG. 2 is a sample flow diagram with color-coded streams: steams—red, water—blue, recycled water—green and wastewater—brown. This approach models the ionic equilibria of the soluble contaminants, providing information about water quality for each unit operation. Modeling a water system also requires creating a “salt” or contaminant balance. The optimal approach is to profile water

deposition and microbiological populations within the cooling water circuit. Quality repurposing. As the quantity and quality of water decreases, industrial users will need to increase their efforts to conserve, recycle water and conform to even stricter regulatory requirements for withdrawal and discharge. Software tools can provide methods for plant personnel to quickly and economically analyze numerous system configurations providing a high level of confidence about the option that best meets their objectives. LITERATURE CITED EPRI document TA-114453, Electric Power Research Institute, Inc., 1999. 2 Data Mobility Systems, a business of Nalco, www. datamobility.com, 2011. 1

LORAINE A. HUCHLER is president of MarTech Systems, Inc., a consulting firm that provides technical advisory services to manage risk and optimize energy- and water-related systems including steam, cooling and wastewater in refineries and petrochemical plants. She holds a BS degree in chemical engineering, along with professional engineering licenses in New Jersey and Maryland, and is a certified management consultant.

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Water sources

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Pretreatment

DWSD

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102 OCTOBER 2012 | HydrocarbonProcessing.com

Waste treatment

West plant

Internal sink

FIG. 1. Sample water-pinch diagram.1

Process/ letdown

550/150/50 psig

City water

Water demands Water pinch Wastewater

quality throughout the system by analyzing numerous samples at every location, and then compare the actual water quality to the predicted value in the model. When the actual water quality and flowrates closely match the predicted values, the model is considered “validated” for the present plant conditions—the baseline case. A validated model allows the facility to hypothetically reprocess water to meet the specification limits for individual process units and to identify candidate water streams for reuse or retreatment (recycling). The model also provides insight into the hydraulic and chemical impacts on the unit and the total system balances for mass and salt concentrations. This computerized modeling provides an accurate assessment of options: different configurations and/or operating scenarios to improve system operability, justification capital improvement projects, optimization system reliability and minimization of the risk of off-spec or lost production. Embedded within this analysis is a projection of the chemistry change for the cooling water. The only remaining analysis for this unit operation is a separate modeling task using a different software program to design an appropriate chemical-treatment program to control the corrosion,

FIG. 2. Refinery water flow schematic.2

Process waste

Process/ letdown Sludge storage

Linde Process Plants, Inc. Accepting Challenges. Creating Solutions.

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A member of The Linde Group Linde Process Plants, Inc. 6100 South Yale Avenue, Suite 1200, Tulsa, Oklahoma 74136, USA Phone: +1.918.477.1200, Fax: +1.918.477.1100, www.LPPUSA.com, e-mail: sales@LPPUSA.com Linde Process Plants, Inc. is an Equal Employment Opportunity Employer

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