Hydrocarbon Processing 01 2012

November 1, 2017 | Author: Ionela Poenaru | Category: Natural Gas, Liquefied Natural Gas, Oil Refinery, Algae Fuel, Petroleum
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A radically different transport sector


Update on replacing large columns in revamps

Scaling up renewables

Innovative technologies treat shale gas, improve LNG operations and more

Are you losing money with your controllers?


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JANUARY 2012 • VOL. 91 NO. 1 www.HydrocarbonProcessing.com


41 45 49 55

Viewpoint Prominent executives and analysts from the natural gas sector share their insights on market trends and future opportunities for gas, including the development of shale gas resources, new liquefied natural gas (LNG) applications, the changing landscape for LNG trade, growth in gas-fired power generation, and more.

Overcome challenges in treating shale gases Manipulating process plant parameters helps meet pipeline specifications R. H. Weiland and N. A. Hatcher

Innovative APC boosts LNG train production APC application yields significant operability, economic benefits A. Taylor and S. Jamaludin

Apply new enhanced tubes to optimize heat transfer in LNG trains New developments for heat exchangers reduce capital and plot size of key equipment B. Ploix and T. Lang

61 65


A radically different transport sector by 2050


China dominates the nylon engineering plastics market


New policies needed to scale up renewable energy


Bioplastics demand to exceed 1 million metric tons in 2015


Cloudy outlook

Select optimal schemes for gas processing plants Careful process evaluation helps meet product requirements and environmental standards M. Maleki and M. Khorsand Movaghar

Improve process control for natural gas heat exchangers Dynamic simulation model identifies how to optimize plant controllability and safety H-M. Lai



Cover The Karratha Gas Plant, located north of Perth, Australia, is one of the world’s most advanced and integrated gas production systems. The facility produces LNG from five trains, domestic gas from two trains, condensate from six trains and LPG from three fractionation units. Photo courtesy of Woodside.

Consider lobe blowers combined with compressors New blower meets low-pressure applications cost-effectively H. P. Bloch



Are you losing money when tuning controllers? Here are 10 rules, if followed, that will result in poor process performance M. J. King




HPINSIGHT Global HPI: 90+ years old and still going strong


HPIN RELIABILITY Dealing with asset management and life extension


HPINTEGRATION STRATEGIES Inline blending can help process plants cut costs and reduce quality give-away


HPIN ASSOCIATIONS Making safety second nature


ENGINEERING CASE HISTORIES Case 66: Fiberglass mixing tank flexing vibration

How to manage vaporization in an analytical system When done properly, this process ensures that all compounds vaporize at the same time, preserving the sample’s composition D. Nordstrom and T. Waters



Case history: Replacement of existing pressure vessel Installing new equipment involves more processes to ensure safety and to meet new codes D. Fearn and J. McKay


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

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

For more information about article reprints, call Rhonda Brown with Foster Printing Company at +1 (866) 879-9144 ext 194 or e-mail [email protected] HYDROCARBON PROCESSING (ISSN 0018-8190) is published monthly by Gulf Publishing Company, 2 Greenway Plaza, Suite 1020, Houston, Texas 77046. Periodicals postage paid at Houston, Texas, and at additional mailing office. POSTMASTER: Send address changes to Hydrocarbon Processing, P.O. Box 2608, Houston, Texas 77252. Copyright © 2012 by Gulf Publishing Co. All rights reserved. Permission is granted by the copyright owner to libraries and others registered with the Copyright Clearance Center (CCC) to photocopy any articles herein for the base fee of $3 per copy per page. Payment should be sent directly to the CCC, 21 Congress St., Salem, Mass. 01970. Copying for other than personal or internal reference use without express permission is prohibited. Requests for special permission or bulk orders should be addressed to the Editor. ISSN 0018-8190/01.

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I JANUARY 2012 HydrocarbonProcessing.com

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5 μS/cm) and is ideally suited for applications in the food and beverage, water and wastewater, and other process industries. The Promag 53 flowmeter features an integrated web server that allows authorized users to remotely view flow data, conduct diagnostics, configure the flowmeter and perform process optimization. Data can also be securely accessed by higher-level software such as enterprise resource planning (ERP) systems, process historians, control-loop tuning programs, and asset management systems. By using EtherNet/IP, up to 10 variables can be configured, including volume flow, calculated mass flow and totalized flow for remote access. Traditionally, devices measuring and controlling process variables rely on a process instrumentation network to transfer

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AdvaSulf™ Claus Tail Gas Treatment, Clauspol® II, Sulfreen™, Sultimate™ I Sulphur degassing, Aquisulf™



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FIG. 3

The electromagnetic flowmeter features EtherNet/IP connectivity for integration with process automation.

HPINNOVATIONS data, while other devices within the plant work on a completely different network. By improving this complex, multi-tier networking strategy with one standard network architecture, namely EtherNet/ IP, users have better access to real-time information. This improves the ability to monitor overall performance, troubleshoot out-of-margin conditions and minimize downtime. In addition to EtherNet/IP connectivity, the Promag 53 has built-in connectivity to FOUNDATION fieldbus, PROFIBUS, MODBUS and HART. Promag 53 is designed to measure most liquids with a minimum conductivity of 5 μS/cm, and flowrates up to 1,250 gal/min. A minimum conductivity of 20 μS/cm is required for measuring demineralized water. The flowmeter operates in temperatures of –4°F to 140°F (–20°C to 60°C) and pressures up to 580 psi. Select 5 at www.HydrocarbonProcessing.com/RS

Gas-treating simulation tool offers ammonia calculations Optimized Gas Treating Inc. has released Version 5.0 of ProTreat, its gas-treating process-simulation tool, which includes the option to add ammonia as a component for which absorption and stripping are calculated on the basis of mass transfer rates. This addition enables users to model sour water strippers and to determine the impact of ammonia as a contaminant in amine systems. Another feature of ProTreat Version 5.0 is enhanced reporting of stream data in order to provide phase-specific compositions; thermodynamic, physical and transport properties for use in generating heat exchanger curves; and for other engineering tasks. Improved methane solubility predictions in methyldiethanolamine (MDEA), based on recent university research data, have also been incorporated. The new ammonia package will be particularly useful for design, optimization and troubleshooting in refinery and syngas applications. Select 6 at www.HydrocarbonProcessing.com/RS

Bentley expands software for Microsoft Through its commitment to the Microsoft Azure Platform partner program, software solutions company Bentley Systems Inc. recently expanded its strategic relationship with Microsoft Corp. Bentley is bringing a broad range of Azure-cloudbased services for sustaining infrastructure

to architecture, engineering, construction and operations (AECO) worldwide at an accelerated pace. Initial offerings on Azure will include the new Bentley Transmittal Services (BTS), enabling AECO organizations to accurately and securely package, deliver, receive and track transmittals through a dashboard portal. These shared services will benefit the users of both the ProjectWise collaboration platform and the AssetWise platform for operations information modeling by reducing risk, saving time and providing greater visibility into project status. BTS includes a dashboard that provides notifications and links to a secure transmittal portal where organizations can see all the transmittals that pertain to their projects. Users will have the option to deploy the portal onsite or online and connect with their existing ProjectWise or AssetWise implementations. The same dashboard serves as a transmittal registry, recording all acknowledgments and tracking all status changes. BTS will include creation, publishing, delivery, response, tracking and status functionalities. Bentley Transmittal Services is currently available onsite with AssetWise. BTS for ProjectWise is available onsite through Bentley’s early adopter program and will be commercially released in Q1 2012. In addition, BTS for the Microsoft Azure platform will be available online in 2012. Select 7 at www.HydrocarbonProcessing.com/RS

Invensys revamps SimSci-Esscor ROMeo optimization software Invensys Operations Management recently released Version 6.0 of its SimSciEsscor ROMeo optimization software (Fig. 4). The most recent version incorporates four new refinery process models that simulate and optimize reforming, coking, isomerization and visbreaking units. The software also contains several new capabilities, including the ability to openly share information using the object linking and embedding for process control— unified architecture standard (OPC-UA). The OPC-UA standard allows the ROMeo software to communicate with many of the company’s simulation and workforceenablement offerings, including its DYNSIM, PRO/II and ArchestrA Workflow software, as well as any third-party product that also uses the OPC-UA standard. ROMeo facilitates equipment monitoring,

utilities optimization and material balance in open- or closed-loop mode. Harpreet Gulati, director of design and optimization at Invensys, noted that the software allows refiners to improve crude selection, evaluate crude supply and reliably predict refinery yields and qualities. It also helps determine the potential for improving yields of higher-value products. Additionally, the integration of ROMeo with Invensys’ Wonderware Intelligence software assists plant personnel in making decisions that reduce operating costs, increase throughput and maximize profit. Select 8 at www.HydrocarbonProcessing.com/RS

ExxonMobil expands energyefficient industrial lube line ExxonMobil Lubricants and Petroleum Specialties Co., a division of ExxonMobil Corp., has added two Mobil SHC highperformance synthetic oils to its industrial lubricants line. The upgraded Mobil SHC 600 Series high-performance synthetic circulating and gear lubricants family is recommended for use in 1,800 applications by more than 500 major equipment builders, and is ideal for use in a wide range of industrial applications. The Mobil SHC Gear Series offers fully synthetic, industrial gear oils qualified by major gear original equipment manufacturers (OEMs) to meet the latest requirements. Developed through extensive research and testing with leading OEMs, Mobil SHC 600 lubricants and Mobil SHC Gear Series are formulated to deliver energy-efficiency savings of up to 3.6% compared to conventional oils (when tested in a worm gearbox under controlled conditions) and to optimize the performance of equipment operating in extreme conditions. In addition to the energy-efficiency benefits, the new Mobil SHC lubricants offer a service life of up to six times longer than competing mineral oil-based gear lubricants. Select 9 at www.HydrocarbonProcessing.com/RS

FIG. 4

The optimization software provides process models for the refining, petrochemical and gas processing industries.


I 33

HPI MARKET DATA 2012 YOUR GUIDE TO PROFITABLE PLANNING IN 2012 AND BEYOND Produced by the staff of Hydrocarbon Processing, HPI Market Data 2012 is the industry’s most trusted forecast of capital, maintenance and operating expenditures for the petrochemical, refining and natural gas/LNG industries. Order your copy and gain actionable insight and analysis to drive your planning and global activities towards increased profitability and market share in 2012 and beyond.

Order Online at GulfPub.com/2012HPI or Call +1 (713) 520-4426 Strategic Planning • Market Analysis and Trends • New Growth Opportunities


North America BioAmber and Mitsui & Co. have partnered to build and operate a manufacturing facility in Sarnia, Ontario, Canada. The initial phase of the facility is expected to produce 17,000 metric tons of biosuccinic acid, and commercial production is anticipated in 2013. The partners intend to expand capacity and produce 35,000 metric tons of succinic acid and 23,000 metric tons of 1,4 butanediol (BDO) on the site. Bioamber and Mitsui also intend to jointly build and operate two additional facilities. These facilities, together with Sarnia, will have a total cumulative capacity of 165,000 tons of succinic acid and 123,000 tons of BDO. BioAmber will be the majority shareholder in the plants. Additionally, the partners plan to build and operate a second plant in Thailand, which is projected to come online in 2014. They are undertaking a feasibility study for the Thailand plant with PTT MCC Biochem Co. Ltd., a joint venture established between Mitsubishi Chemical Corp. and PTT Public Co., Ltd. BioAmber and Mitsui & Co. also plan to build and operate a third plant, located in either North America or Brazil, that will be similar in size to the Thailand project. SNC-Lavalin has a major contract from an oil-sands mining producer to provide engineering, procurement and construction (EPC) services for a froth-treatment plant in the Fort McMurray region of Canada. The contract value is in excess of $650 million. The froth-treatment plant will process 155,000 bpd of feedstock from the bitumen extraction plant in the form of bitumen froth. The engineering phase is now underway and construction is scheduled to begin in February 2012. Mechanical completion for the construction is expected in September 2014. GT Logistics, LLC (GTL) has began installing rail lines at its OmniPort location in Port Arthur, Texas. The OmniPort is expected to open for business in January 2012, serving as a multimodal terminal for crude oil and other products transported via rail, ship, barge and truck.

The $95 million, 1,100-acre facility neighbors refineries with over 1 million bpd of capacity, and multiple chemical and processing plants, and is located less than one mile from over 4 million bbl of petroleum product-storage capacity and pipelines serving the region. The rail terminal, served by Union Pacific, will be able to receive unit train traffic, with 300 acres of rail-car storage onsite that will be capable of storing, switching and transloading over 1,000 rail cars. The rail terminal site also features a multibarge receiving dock on Taylor’s Bayou; convenient access to Highway 73 and Interstate 10; and connectivity to the region’s extensive network of pipelines. The initial phase of the rail, drainage and road improvement construction began earlier in the year and will be completed by the end of 2011. Cheniere Energy Partners, L.P. has selected Bechtel to provide engineering, procurement and construction (EPC) services for two new liquefaction trains at the Sabine Pass liquefied natural gas (LNG) terminal in Cameron Parish, Louisiana. The project builds on Bechtel’s previous work at Sabine Pass, where the company designed, built and expanded the LNG receiving facility. Bechtel will design, construct and commission the two liquefaction trains using ConocoPhillips’ Optimized Cascade technology. The liquefaction trains will be built next to the existing facilities at the Sabine Pass LNG terminal, which include five tanks with storage capacity of 16.9 billion ft3 equivalent, two docks that can handle vessels up to 265,000 m3 and vaporizers with regasification capacity of 4.0 billion cfd. Construction is expected to begin in 2012.

ing waste heat generated at compressor stations along the Alliance Pipeline system to produce emission-free electric power. The company has four waste-heat recovery units operational at Kerrobert, Loreburn, Estlin and Alameda, Saskatchewan. While its Whitecourt Recovered Energy Project (WREP) marks the company’s fifth waste-heat recovery installation, it is the first to use GE’s ORegen system. Construction of the WREP will commence in May 2012.

South America MODEC, Inc., has commissioned a UOP Separex membrane system for processing natural gas on a new floating production, storage and offloading (FPSO) vessel. The FPSO is using the Honeywell UOP Separex membrane system and adsorbents to remove carbon dioxide and water from 5 million standard m3/day of natural gas from the Lula oil field off the coast of Brazil. The FPSO was commissioned in July 2011. A second FPSO, still in construction and also using UOP Separex technology, is expected to be commissioned in September 2012. The Lula oil field is said to contain the largest oil discovery in the Western Hemisphere in the last 30 years and is believed to contain 8.3 billion bbl of oil and natural

Trend analysis forecasting Hydrocarbon Processing maintains an extensive database of historical HPI project information. The Boxscore Database is a 35-year compilation of projects by type, operating company, licensor, engineering/constructor, location, etc. Many companies use the historical data for trending or sales forecasting. The historical information is available in comma-delimited or Excel® and can be custom

NRGreen Power and GE have plans for a new recovered energy project that will produce power without additional emissions using the first global application of GE’s innovative ORegen system. The technology will be installed at Alliance Pipeline’s Windfall Compressor Station near Whitecourt, Alberta, Canada, to generate electricity through the use of waste heat. NRGreen Power Ltd. Partnership works to develop clean energy by convert-

sorted to suit your needs. The cost depends on the size and complexity of the sort requested. You can focus on a narrow request, such as the history of a particular type of project, or you can obtain the entire 35-year Boxscore database or portions thereof. Simply send a clear description of the data needed and receive a prompt cost quotation. Contact: Lee Nichols P.O. Box 2608, Houston, Texas, 77252-2608 713-525-4626 • [email protected] HYDROCARBON PROCESSING JANUARY 2012

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HPIN CONSTRUCTION gas. The field is operated by Petrobras in partnership with BG and Galp. Honeywell’s UOP Separex technology upgrades natural gas streams by removing carbon dioxide and water vapor. These contaminants must be removed to meet the quality standards specified by pipeline transmission and distribution companies, as well as end users of the natural gas. Haldor Topsøe has signed an agreement with Petrobras for the supply of critical equipment and materials for two SNOX plants. The plants will be installed at the new RNEST grassroots refinery in Pernambuco, Brazil. The supply covers internals for 80 wet-gas sulfuric-acid (WSA) condensers for condensation of sulfuric acid, eight units for acid mist control and a complete acid system. In addition to treating the boiler flue gases, the SNOX plants will also treat Claus plant tail gases, amine gases containing hydrogen sulfide and sour-water stripper (SWS) gases containing ammonia. These SNOX plants are designed for the possible elimination of the Claus plants, which then means that all the refinery’s sulfur compounds are converted into sulfuric acid. The two SNOX plants will be installed in parallel and will each treat up to 650,000 Nm3/h of flue gas while producing up to 750 metric tpd of sulfuric acid. In addition to producing sulfuric acid, the SNOX plants will also export up to 100 ton/hour of high-pressure steam to the refinery steam grid. The contract for basic engineering was signed with Petrobras earlier and has already been executed. The supply of equipment will take place during the coming 16 months, and startup of the SNOX plants is planned for 2013.

Europe Jacobs Engineering Group Inc. has a contract to provide detailed engineering, procurement support and support services during construction of a new ester production plant at OXEA GmbH’s existing manufacturing facility in Oberhausen, Germany. The Esterplant 2 project is part of OXEA’s strategy to expand global ester production capacity by 40% to meet the growing global demand for OXEA’s esters. These specialty chemical products are replacing the traditional phthalate plasticizers. The fast-track project is expected to come onstream in 2012. 36

I JANUARY 2012 HydrocarbonProcessing.com

MAN Diesel & Turbo is installing a CHP cogeneration unit at the Rheinberg production plant of Solvin GmbH & Co. KG, a joint enterprise of Solvay and BASF. The first of MAN’s new 6-MW gas turbines will be used commercially for the plant, which manufactures chemical products including polyvinyl chloride (PVC). The new CHP plant is designed to supply 6 MW of electrical and 11 MW of thermal power, thus enabling Solvin to meet its own electricity requirements in the future. ThyssenKrupp EPC contractors have supplied and commissioned a plant for producing 3,500 tpd of urea solution in Sluiskil, Netherlands, for Yara of Norway. The plant, which took three years to build, has now been handed over to the customer. Yara invested €400 million in its construction. The plant meets the latest environmental standards, as well as the best available technology standards. There are even special collection systems that, should the plant malfunction, ensure that no hazardous substances escape into the environment. What really makes the plant so remarkable, though, is that some of the urea it produces will not be used as fertilizer but as an aqueous urea solution to treat diesel exhaust fumes. This technology, known as AdBlue, reduces NOx emissions. The technology was licensed by Stamicarbon. ThyssenKrupp Uhde was responsible for the engineering, equipment supply and plant construction on a fixed-price, turnkey basis. Neste Oil is building a system for recovering emissions released when loading ships at the harbor of its Porvoo refinery. The system, valued at approximately €23 million, will recover the majority of the volatile organic compounds (VOCs) released into the atmosphere when loading gasoline. The new system will reabsorb VOCs into gasoline during loading with the help of two absorption tanks and related equipment at the harbor, after which the gasoline used will be returned to the refinery for re-use. A similar system is already in use when loading tanker trucks at the Porvoo refinery’s distribution terminal. Construction work on the VOC recovery system began in October 2011 and the facility is due to be commissioned in the latter half of 2013.

Africa Kenya Petroleum Refineries Ltd. (KPRL) has implemented a new solution from IBM to increase the productivity and efficiency of the company’s oil refinery operations in East Africa. The agreement was finalized by IBM’s business partners Computer Source Point Ltd. and Powertech IST Data. The new IBM solution will allow KPRL to manage, measure and track the life cycle of its oil-processing equipment such as pipes, heat exchangers, pumps, valves, boilers, furnaces, compressors, tanks and turbines. Niger Delta Petroleum Resources Ltd. (NDPR), the fully owned subsidiary of Niger Delta Exploration & Production Plc (NDEP), has been granted a license to operate (LTO) the Ogbele mini refinery. This license, granted by the federal government of Nigeria, is said to be the first of its kind to be granted to an independent, publicly owned Nigerian company. It gives NDPR full authority to operate its mini diesel refinery (topping plant) at the Ogbele Field in old OML 54 (Rivers State). The LTO will make NDPR’s mini diesel refinery the first independently owned and fully operational diesel refinery in Nigeria. The refinery has an initial capacity of 1,000 bpd. It commenced production in December 2010, using crude-oil feed from NDPR’s existing Ogbele flowstation.

Middle East Tecnimont S.p.A., the main operating company of Maire Tecnimont S.p.A., has an engineering procurement, construction and commissioning (EPCC) contract on a lump-sum turnkey basis for a new fertilizers complex within the existing industrial area in the Aswan Governorship in Upper Egypt, from the Egyptian Chemical & Fertilizers Industries–KIMA. The fertilizers complex will comprise an 1,200-tpd-capacity ammonia-production unit, implementing KBR’s Purifier technology; one 1,575-tpd-capacity ureamelt production unit, implementing Stamicarbon’s Pool Reactor technology; one 1,575-tpd-capacity urea-granulation unit, implementing Stamicarbon’s urea-granulation technology; and all the necessary utilities and offsite facilities to support the process units. The overall project value is approximately $540 million and completion is expected by the end of July 2014.

HPIN CONSTRUCTION Toyo Engineering Corp. was awarded an energy optimization project for one of SABIC’s existing ammonia plant and package boilers at the Al-Jubail Fertilizer Co. (Al-Bayroni) in the eastern region of the Kingdom of Saudi Arabia. The plant, which has a production capacity of 1,300 metric tpd of ammonia, has been in operation since 1983. Toyo will reduce and optimize the energy consumption in similar ammonia plants and package boilers. Project implementation is scheduled to be completed in the second quarter of 2013. Toyo’s scope of work includes engineering, procurement, construction, pre-commissioning and commissioning assistance on a lump-sum turnkey basis. Qatar National Facilities Services, a Qatari-based company partly owned by Fluor, has signed a five-year comprehensive maintenance-services contract with RasGas in the industrial city of Ras Laffan, Qatar. The contract was awarded to provide maintenance services for the entire complex. Fluor previously completed the RL3 Common Offplot project for RasGas in 2009. For that project, Fluor trained more than 62,000 different workers from 40 different countries at the site, with peak construction manpower reaching nearly 9,000 workers in January 2008.

naces. Equipment delivery will be completed in April 2013. The North refinery in Baiji is said to be one of the largest refineries in the Republic of Iraq, constructed by Chiyoda in 1983 with a capacity of 150,000 bpd. At the end of February 2011, hydrotreater furnaces at the North refinery were shut down by bomb blasts. NRC requested Chiyoda for planning assistance of shortterm emergency measures, and Chiyoda was awarded the contract for the furnacereplacement work as a permanent solution through international tender.

Asia Pacific Davy Process Technology Ltd., a Johnson Matthey company, and The Dow Chemical Co.’s Oxygenated Solvents Business, have announced that Wison (Nanjing) Clean Energy Co., Ltd., has selected LP Oxo SELECTOR 10 technology for its new oxo alcohols plant in Nanjing, China. With this licence, Wison Energy will build a LP Oxo plant with a capacity of 125 kiloton/ yr of 2-ethylhexanol and 125 kiloton/yr of butanols.

Wison Energy’s Nanjing plant operates a 600-kiloton/yr carbon-monoxide plant and supplies carbon-monoxide, synthesis gas, hydrogen and methanol to other facilities located in the Nanjing Chemical Industry Park. Michelin Siam Co., Ltd., has awarded Technip a lump-sum turnkey engineering, procurement and construction (EPC) contract for a new elastomer composite plant to be built in Southern Region Industrial Estate, Songkhla Province, Thailand. The contract is in line with Technip’s strategy to expand its business base, including its onshore segment. It covers preliminary engineering, detailed engineering, project management, procurement, construction, pre-commissioning and commissioning, and startup assistance. The plant will produce rubber composites. Technip’s operating center in Bangkok, Thailand, will execute the contract, which is scheduled to be completed at the beginning of 2013. HP

Cellier Activity of ABB France’s Process Automation Division has started up a new lube-oil blending plant (LOBP) for Petromin Corp. in Jeddah, Saudi Arabia. Cellier Activity was responsible for the detailed engineering and procurement, mechanical and electrical site supervision, along with commissioning of the core process equipment. The scope of supply included both batch- and inlineblending technologies. The new facility is automated, and plant-wide activities are managed by a Lubcel control system to ensure process flexibility and safety. With an annual capacity of 125,000 tons of lubricating oils per shift, the new facility is able to produce a wide range of automotive and industrial oils, and is reportedly one of the largest LOBPs in the Middle East. North Refineries Co. (NRC) has awarded Chiyoda Corp. a contract for replacement of furnaces at the North refinery in Baiji, Republic of Iraq. The contract covers engineering, procurement and delivery of three sets of furSelect 154 at www.HydrocarbonProcessing.com/RS





Ex Capacity Unit

Egypt Morocco Nigeria


Ammonia DAP (4) Refinery


1200 tpd None 100 bpd

Senegal Zambia

PCMC Indeni Refining Co.

Aswan Jorf Lasfar Imo State, Egbema/ Oguta Industrial Park Dakar Ndola

Refinery Refinery


100 bpd 24 bpd


Curtis Island


Sinopec IOCL Shell Hazira/Total JV Indian Oil Corp Ltd Essar Oil Ltd Amerind Petroleum Pvt Ltd. Pertamina/PT Chandra Asri/ Saudi Aramco JV Michelin Siam Elastomer BioAmber/Mitsu & Co. Petrovietnam

Changling Gujarat Hazira Paradip Vadinar Visakhapatnam Cilegon

Crude Unit Refinery LNG Terminal Refinery Distiller, Crude Refinery Refinery

Songhkla Undisclosed Nghi Son EZ

Elastomers Biosuccinic Acid Refinery

Socar StatoilHydro Rompetrol Rafinare Bashneft Moscow Oil Refinery Novo Ufimskii NPZ Lukoil-Volgograd Neftepererabotk Turkmengas

Baku Mongstad Constanta, Petromidia Refinery Bashkortostan Moscow Ufa Volgograd Undisclosed

Processing, Oil/Gas VOC Recovery Hydrogen Hydrocracker Refinery Hydrotreat, Gas Oil Coker, Delayed Refinery (3)

Qeshm Island Kuwait Jazan Jeddah Ras Tanura Yanbu Ruwais

Refinery, Heavy Ends Olefins Refinery Blending, Lubes Petrochemical Complex Refinery Polypropylene (2)

400 125 400 400 480

Commerce City Mc Pherson Cameron Parish Meraux Pascagoula Tulsa Port Arthur

Biorefinery Coker, Delayed LNG Liquefaction Plant Refinery RE Desalter, Crude RE Electrofining Crude Unit EX

20 400 16.9 135 50 18 600

Cost Status Yr Cmpl Licensor





2014 2015 2016

1200 600





59.2 91 1272 1620 505 7000


2013 2022 2013 2012 2011 2014 2014



2013 2014 2016

300 bpd 1500 36 Msm3/hr 18.7 40 MNm3/h 99 None 551 10 MMtpy 26.4 2 MMtpy 309.5 None None


2020 2011 2012 2016 2020 2012 2011


2013 2011 2016 2011 2015 2014 2013


2011 2015 2012 2011 2012 2011 2012

Jacobs Engineering SA

KBR|Tecnimont Tecnimont Jacobs Engineering SA Jacobs Engineering SA

ConocoPhillips Ltd



Indian Oil ABB Lummus

BOS|Tecnimont FW ABB Lummus

Jacobs Essar





Technip|FW|UOP Aker Solutions Technip


ASIA/PACIFIC Australia Marine|Bechtel China India India India India India Indonesia Thailand Thailand Vietnam

3.8 MMtpy EX EX RE

160 15 3.6 15 14 7.5 300

bpd Mtpy Mtpy MMtpy MMtpy Mtpy bpd

None m-t 200 Mbpd

EUROPE Azerbaijan Norway Romania Russian Federation Russian Federation Russian Federation Russian Federation Turkmenistan



Aker Solutions Technip

Vnipineft Axens

MIDDLE EAST Iran Kuwait Saudi Arabia Saudi Arabia Saudi Arabia Saudi Arabia UAE

Iranian Oil Rfg Kuwait Petro Corp Saudi Aramco Petromin Sadara Chemical Co. Saudi Aramco\ConocoPhillips Borouge III

30 bpd None bpd tpd Mbpd Mbpd Mtpy

7000 20000 1300 722

Dow KBR Tecnimont

Axens|KBR|CLG Axens ABB Cellier Jacobs |KBR ABB|KBR|Jacobs Aramco Services Co|KBR Hyperion|Samsung Eng

UNITED STATES Colorado Kansas Louisiana Louisiana Mississippi Oklahoma Texas

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Purvin & Gertz

Jones Industrial Holdings, Inc.

has been acquired by

$7,000,000 Senior Subordinated Notes

Ventech Project Investments L.P.

financing provided by

IHS Inc. GulfStar Group acted as exclusive financial advisor to Purvin & Gertz Inc.

GulfStar Group acted as exclusive financial advisor to Jones Industrial Holdings, Inc. and arranged the private placement of these securities.

Ventech Engineers L.P.

has completed a majority recapitalization with

$200,000,000 Infrastructure Project Investment Fund GulfStar Group arranged, structured and negotiated this financing and acted as financial advisor to the Ventech companies.

Thorpe Corporation

has completed a majority recapitalization with

Cooper Investment Partners

The CapStreet Group, LLC

GulfStar Group acted as exclusive financial advisor to Ventech Engineers L.P.

GulfStar Group acted as exclusive financial advisor to Thorpe Corporation.

Securities offered through GulfStar Group I, an affiliated entity, member FINRA – SIPC.

Our process begins with a thorough understanding of our client's needs and objectives to successfully engineer transactions refined to achieve their goals. GulfStar is a leading middle market investment bank headquartered in Houston, Texas, and has been named one of the top 10 most active investment banks and advisors in the energy industry over the past decade according to The PitchBook Decade Report. Having closed more than 70 transactions in the energy and manufacturing industry within the past 5 years, GulfStar has the industry experience and transactional leadership to deliver superior results. 0HUJHUV $FTXLVLWLRQV$GYLVRU\6HUYLFHV‡5HFDSLWDOL]DWLRQV‡,QVWLWXWLRQDO3ULYDWH3ODFHPHQWV‡&RUSRUDWH)LQDQFLDO$GYLVRU\

Colt Luedde Managing Director [email protected]

Kent Kahle Managing Director [email protected]

Cliff Atherton Managing Director [email protected]

Steve Lasher Managing Director [email protected]

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HPI VIEWPOINT The opportunities of abundance: How shale gas changes the energy landscape Sara Banaszak, chief economist and vice president of America’s Natural Gas Alliance (ANGA), draws on her 15-year background in natural gas and oil to help guide ANGA’s research and analysis and to develop and promote policies that reflect the potential of abundant new natural gas supplies. Her knowledge of the energy industry ranges across the upstream, downstream, international and domestic arenas, and covers policy, regulatory and commercial issues. She has conducted extensive analysis of global liquid natural gas (LNG) trade, analyzed the economic impacts of US energy policies, modeled the dynamics of oil refineries, and published research on oil demand in Asia. Ms. Banaszak has served as a speaker, chair and organizing committee member for major North American and international conferences, including Gastech. She has also worked in consulting, directing PFC Energy’s North American Gas Policy Service, and previously with FACTS Global Energy. Other professional experience includes research for the American Petroleum Institute and the EastWest Center, project management at the US Department of Energy, and international energy modeling for the US Energy Information Administration (EIA). Additionally, Ms. Banaszak is an active member of several associations, including the National Association of Business Economists and the International Association for Energy Economics. She holds a master’s degree in applied economics from the University of Hawaii and a bachelor’s degree in international relations from the University of Pennsylvania.

In 2005, the US EIA forecast that US demand for natural gas would be met by importing greater volumes of the fuel in the form of LNG. Specifically, net LNG imports were expected to rise to 10% of consumption in 2010 and to 20% in 2025. Six years later, however, the data for 2010 indicate that net LNG imports were less than 2% of US natural gas consumption, while domestic production has far exceeded earlier forecasts. These unanticipated changes are the result of new production of natural gas from shale resources, a phenomenon that is changing the energy landscape. Today, the US has a growing supply of domestic natural gas that can power the country for generations to come. According to the EIA, the total natural gas resource base sits at 2,543 trillion cubic feet (Tcf ), powered by new discoveries across the US. These discoveries have fundamentally transformed the long-term supply outlook. Shale gas production makes up more than 20% of US supply, an increase from 1% in 2000, and energy analysts at ICF International predict that, in about 25 years, this figure will grow to around 65% of total supply. In short, shale gas supplies that were initially described as an “unconventional” resource will become conventional in the near future. The arrival of lower and more stable prices. The

increases in North American natural gas resources and production have brought new stability to gas markets. In fact, record production and low prices were sustained right through the worst

economic downturn since the Great Depression. The US reached record levels of annual gross natural gas production in each of the years from 2008 through 2010. At the same time, natural gas prices (NYMEX front-month contract) have not exceeded $6.10 per million Btu (MMBtu) since January 2009, and in the first three quarters of 2011 prices have averaged $4.21/MMBtu. EIA’s long-term outlook anticipates that prices will stay below $7/MMBtu until at least 2035. There has also been a clear break in the traditional linkage between oil and natural gas prices. In mid-2008, as the price of crude oil peaked at $140 per barrel (bbl), natural gas hit a historic high of $14/MMBtu. After a drop in prices, oil jumped to over $100/bbl and has remained above $75/bbl over the last year (NYMEX front-month contract), while natural gas prices stayed at much lower levels (under $4.50/MMBtu) than would have been anticipated with a closer price linkage. Gas opportunity in the power sector. The realization in the US of abundant natural gas supplies coincides with a growing interest in cleaner sources of energy and improved energy security through the expanded use of domestic fuels; this opportunity is particularly strong in the power sector. In US power generation, natural gas is greatly under-utilized. Gas-powered plants currently make up the greatest portion of US generation capacity, yet only about 25% is actually utilized. Low-priced natural gas supplies, coupled with power plant infrastructure already in place around the US, are already facilitating expanded power-sector use, which increased more than 25% between 2005 and 2010. Continuing to expand powersector use also has an immediate and significant benefit for the environment. In fact, the Congressional Research Service (CRS) found that, if the US doubled the use of natural gas combinedcycle plant capacity, nearly 10% of the country’s CO2 emissions would be displaced—not to mention reductions in pollutants such as mercury, NOX and SOX. Driving change. The transportation sector is also benefit-

ting from newfound natural gas abundance in the US. The gas-producing companies that are part of ANGA have been driving change in many of the same locations where natural gas is produced. These companies are building fueling facilities for natural gas vehicles (NGVs), buying and using NGVs, financing new technologies for NGVs and selling natural gas to other NGV users. The use of natural gas for fleet vehicles has been expanding rapidly. In Los Angeles, California, the local metro system retired its last diesel bus in January 2011 and now has 2,221 buses (99.6%) of its fleet running on compressed natural gas (CNG). The commercial use of NGV fleets is also rising, most notably with companies like UPS and AT&T. Natural gas is also competitive in the heavy-duty sector of the transportation market, as these vehicles travel the most miles and have significantly lower fuel economy and substantial fuel HYDROCARBON PROCESSING JANUARY 2012

I 41

HPI VIEWPOINT needs. Fuel cost is a key driver in making natural gas competitive for transportation. Natural gas costs, on average, one-third less than conventional gasoline. Additionally, NGVs are proven to have lower operating and maintenance costs, both of which generate significant savings over the life of vehicles. Many gaspowered fleets report 15%–28% savings compared to those running on diesel. Shale gas is fueling economic success. The benefits

in the US from a move toward domestic natural gas will also have an important impact on the economy, including lower electricity prices, higher industrial production and higher household incomes. Using more of this domestically sourced fuel will inevitably lead to increased employment for staff geologists, engineers and drilling specialists, as well as an escalation in supporting industries. Shale gas already supports 600,000 jobs, and this number is poised to grow significantly, exceeding 1.6 million jobs by 2035.

Growth in America’s steel and chemical industries is a great example of how the US benefits from the shale boom. The steel industry is seeing increased orders for specialty pipes, while production of ethylene and ethylene-based products is expanding in the US. Moving forward. It is clear that the energy dynamic in the US has changed over the last few years. Importing natural gas was an accepted truth, but today’s reality is that the US has enough natural gas to meet growing demand, allowing the country to focus on domestic uses and export possibilities. The production of shale gas will continue to reverberate throughout global energy markets as planned US imports are displaced and as production of shale gas is extended internationally. In the US, the production of shale gas will continue to help the country move away from foreign fuels, saving precious dollars in the process and addressing some pressing environmental challenges. It is a monumental opportunity that our industry is excited to tackle. HP

Cautious optimism for growth in the natural gas market Edward Kelly, the vice president of North American gas and power at Wood Mackenzie, is a recognized expert analyst of the North American natural gas and power industries. He has focused expertise in the natural gas midstream, energy markets, regulatory issues and energy business strategies. Mr. Kelly leads major consulting engagements for Wood Mackenzie’s Americas Gas and Power consulting group, contributes to the research product, and integrates insights from Wood Mackenzie’s World Oil, North American Gas and Power, and Global LNG practices for the benefit of Wood Mackenzie clients. In the course of his career, Mr. Kelly has advised many North American and international energy companies on business strategy for the North American energy market, and on the effects of global energy market forces on the North American energy industry and marketplace. His advisory experience also includes numerous state governmental and regulatory bodies, as well as utilities, producers, and gas pipeline and storage companies. Prior to joining Wood Mackenzie, Mr. Kelly was director of research for the North American natural gas unit of Cambridge Energy Research Associates. Previously, he worked as an analyst for Panhandle Energy (now Spectra Energy) and Tennessee Gas Transmission. Mr. Kelly holds a BA degree in economics from the University of Chicago and an MBA degree in finance from the University of Texas at Austin.

The natural gas industry is more optimistic now than it has been in perhaps decades, as the fuel is on its way to capturing a growing and significantly larger share of the US energy marketplace. With the advantages of a large resource base, relatively reasonable costs, and environmental pressures on competing fuels, such optimism is justified; however, the industry’s hopes are not guaranteed to be fulfilled. While the resource base itself is less and less in doubt, the extent to which production costs will continue to decline remains uncer42

I JANUARY 2012 HydrocarbonProcessing.com

tain, and regulatory and environmental costs are far from resolved. The demand side faces even more uncertainty, with market growth more dependent than ever on the overall health of the economy and on the increasingly contentious politics of power generation. Natural gas’ share of energy consumption in the US has varied within the 20%–30% range since the early 1970s, down from just over 30% in the late 1960s when natural gas prices were heavily regulated at low levels, heavy industry was running strong, and supply was not an issue. Successive recessions in 1973–1974 and 1981–1982 hit the industrial gas market hard, with the result being that, by 1986, natural gas’ market share in US energy consumption had fallen to less than 22%, and the overall US gas market had dropped to just over 16 trillion cubic feet (Tcf )—the lowest market share for gas since 1953. Recovery was initially headed by the industrial sector during the 1990s “gas bubble,” as a strong economy and cheap natural gas drove a rebound of approximately 8 billion cubic feet per day (Bcfd) in industrial consumption, while power consumption grew by 4–5 Bcfd over the same period. Gas’ share of US energy consumption fell again through the mid-2000s, however, as high, volatile gas prices discouraged industrial consumption, while gasfired power generation continued to show gains. In recent years, gas’ share has increased slowly, climbing back to 25% in 2010 and 2011. This growth is a result of stabilization in the industrial market and continued growth in gas-fired power generation, which is in part due to the displacement of traditional coal markets in the eastern US. Is the gas industry indeed poised for a new era of growth as it captures larger shares of the industrial and power-generation markets? Again, supply is not an issue for a decade or more, at least. Wood Mackenzie estimates that production in nine active shale plays (Marcellus, Haynesville, Barnett, Eagle Ford, Fayetteville, Woodford, Horn River, Montney and Duvernay) alone is capable of growing by more than 20 Bcfd on net from 2011–2020—greater than expected US demand growth—at Henry Hub prices of between $4.00 per million British thermal

HPI VIEWPOINT unit (MMBtu) and $5.50/MMBtu. This forecast holds in spite of an expected doubling of pipeline gas exports to Mexico and nearly 2.5 Bcfd of LNG exports. Of course, there are other shales under development (Utica, Niobrara) as well, and several plays waiting in the wings should prices rise. As a result, no pure exploration success is required to satisfy the needs of the North American gas market, even for the next 20 years. We already have a good idea where the gas we may need is located, and we know it can be produced at less than $6/ MMBtu, even accounting for some increase in costs associated with tightening environmental standards. On the demand side, the near-term picture is less clear. A weak economy and growing generation from renewable sources are dampening the underlying need for gas generation, while an increase in gas prices would cause gas to lose market share to coal for a time. However, politics are coming into play again, with US Environmental Protection Agency (EPA) regulations expected to result in the retirement of between 50 GW and 60 GW of US coal-fired power facilities within the next 5–7 years. If the economy resumes a steady pace of growth, the gas market should begin to shift from a dependence on pushing supplies at low cost into traditional coal markets, toward supplying growing power-generation needs while some existing coal plants

are retired. Exports and continued coal displacement are likely to remain features of the gas market at prices still attractive to US consumers. Of course, the opportunities for growth are more varied. Petrochemical producers, fertilizer manufacturers, steel producers and other industrials are already making or considering investments to take advantage of low gas prices, and Wood Mackenzie expects industrial consumption to rebound, exceeding the 20-Bcfd level again (up from over 18 Bcfd currently) within 5–7 years. The wide spread between natural gas and oil—which is likely to remain wide for the foreseeable future—is attracting ongoing investments in CNG and LNG vehicles and refueling infrastructures as well, although it will be many years before these make a material difference in the overall marketplace. In short, if steady economic growth resumes, the gas industry is poised to supply an increasing share of the nation’s energy needs once again. Will the industry ever get back to a 30% share, as seen in the late 1960s? Power consumption efficiencies, renewable energy generation, and a slower pace of economic growth create headwinds, so the 30% level will probably not be reached until 2025 or 2030. However, the gas market at that time, forecast at nearly 90 Bcfd, could be one-third larger than it is today— not a bad opportunity for those in the business. HP

LNG’s future: Branching out Susan Tucker Walther, Mustang’s LNG engineering manager, is an accomplished engineer with more than 25 years of diverse industry and engineering experience. A registered professional engineer in Texas, she earned her chemical engineering degree with honors from Texas A&M University and began her career at Shell’s chemical/refinery complex in Deer Park, Texas. She has spent the majority of her career in engineering and construction—designing refineries, chemical facilities, LNG regasification terminals and LNG liquefaction facilities—and has led projects in all phases of design and execution. She joined Mustang Engineering in 2007 and has focused exclusively on LNG-related projects in her current role. Mrs. Walther is the author of several publications related to the LNG industry and has spoken at several conferences, including Gastech 2011 in Amsterdam.

The liquefied natural gas (LNG) industry is mature and poised to branch out in many new directions in the coming decades as it continues to grow in capacity. One forecast takes today’s 325 million metric tons per year (metric MMtpy) of installed or under-construction liquefaction capacity to approximately 500 metric MMtpy by 2030. Traditional onshore baseload facilities, particularly in Australia and Africa, will account for the largest part of the 175-metric MMtpy growth during that period. Nontraditional sizes and locations. While most of the

capacity increase will be driven by traditional baseload plants, expect to see substantial growth in LNG facilities in nontraditional locations and sizes, as well. Arguably, the most interesting future location is offshore. The wide field of floating LNG

(FLNG) players seen five years ago has narrowed substantially and is now led by Shell’s Prelude facility, which is pushing the current boundaries of technology and is positioned to become the world’s first operating FLNG later this decade. The industry is watching closely. Once the concept and its many challenges are proven, expect a “rush to be second” and a concerted push to find ways to make the next generation of FLNGs more economical. Also expect a new breed of mid-scale (0.5–2.0 metric MMtpy) liquefaction facilities to come into their own and help monetize the large number of smaller gas fields, particularly offshore and in remote onshore locations. Unlike baseload LNG facilities, where innovation is driven by the need for higher efficiency and larger train sizes, many mid-scale liquefaction processes and their associated pretreatment processes will focus on operational simplicity and minimization of capital expenditures (CAPEX). At the smallest end of the capacity range will be the small-scale (peak-shaver) to micro-scale (vehicle-fuel) liquefaction facilities that will pop up wherever they can best take advantage of local gas availability and LNG markets. In particular, processes aimed at the vehicle fuel market will increasingly be offered in “catalog” sizes of prepackaged equipment. The lowest CAPEX and shortest construction schedules will rule this market. Don’t forget regas. While less glamorous than liquefac-

tion facilities, import/regasification terminals will also see steady growth. Historically, worldwide regas capacity has been about 2–3 times that of worldwide liquefaction capacity. We can expect to see 45–70 billion cubic feet per day (Bcfd) of new regas capacity by 2030. As with liquefaction, most of the growth in regas will be in traditional, baseload-sized terminals. While liquefaction facilities face the challenge of being located “where the gas is” (i.e., increasingly remote or otherwise challenging locations), regas terminals face a rather different challenge of HYDROCARBON PROCESSING JANUARY 2012

I 43

HPI VIEWPOINT being located “where the demand is.” By definition, this tends to be at or near heavily populated, congested ports. With a lack of available land near shore and a strong not-in-my-backyard mentality in many locations, there is already a successful and growing market for floating storage and regasification units (FSRUs). The full-sized FSRU field is dominated by a few proven players. There is another change in the air surrounding many onshore regas terminals. Open-rack vaporizers, a long-time industry standard, are under increased scrutiny for their real or perceived impact on water temperatures and marine life. Submerged combustion vaporizers or other vaporizers using a fired heat source are far from ideal substitutes, due to their fuel consumption and corresponding air emissions. As such, expect increased interest in utilizing other vaporization heat sources such as waste heat and ambient air, or utilizing the cold duty from LNG vaporization to improve the performance of adjacent processes including electrical power generation, air separation, ethane/propane recovery and seawater desalinization. In many regions of the world, LNG import terminals will be closely associated with new power-generation projects to help meet the demand for reliable power. Some terminals will also distribute LNG locally, via truck or train, for vehicle fuel or other domestic uses. A paradigm shift in shipping. For years, LNG carrier designs have moved in primarily one direction: bigger. From 145,000 m3 a decade ago, the newest carriers now transport up to 260,000 m3 of LNG. While this is a good solution for the

large baseload LNG routes, it presents a significant challenge to tomorrow’s small- to mid-scale liquefaction and regas facilities, where the economics cannot support the full-sized LNG storage tanks, jetties and deep draft required for infrequent visits by full-sized carriers. Watch for LNG mini-carriers and articulated barges—currently used in only a few locations such as Scandinavia—to step up and fill what is now something of a void. These carriers will be best used over relatively small distances to move cargoes in the 5,000–30,000 m3 range from small-scale liquefaction sites to small-scale regas or LNG-to-power facilities. The onshore storage size and cost on both ends of this supply chain are reduced, along with the jetty and water depth requirements. Stay focused on safety. To date, the LNG industry has

an outstanding safety record. We cannot become complacent! Nearly all of these described changes involve taking LNG into more remote areas with typically smaller, less LNG-experienced operators and owners. The trends toward mid-, small- and micro-scale LNG will result in an increased number of facilities, albeit smaller than today’s baseload plants and terminals. As an industry, we have to remain diligent to ensure that our designs, equipment, procedures and systems are safe and environmentally sound. LNG’s future is solid and growing, and it will move in exciting new directions during the next few years. There will be many ways to be a part of the growth, so get innovative—and, above all, stay safe. HP

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Overcome challenges in treating shale gases Manipulating process plant parameters helps meet pipeline specifications R. H. WEILAND and N. A. HATCHER, Optimized Gas Treating Inc., Houston, Texas


hale represents an astonishingly large, new source of natural gas and natural gas liquids (NGLs). However, a common misconception seems to be that, for the most part, shale gases are sweet and do not need to be treated. Although not highly sour in the traditional sense of having high H2S content, and with considerable variation from play to play and even from well to well within the same play, shale gas often contains tens or hundreds of parts per million of H2S, with wide variability in CO2. Gas in the Barnett shale play of North Texas, for example, contains several hundred parts per million by volume (ppmv) of H2S and several percentages of CO2—far from pipeline quality. In other shales, such as Haynesville and the Eagleville field of the Eagle Ford play, H2S is known to be present. In other cases, such as the Antrim and New Albany plays, underlying sour Devonian formations may communicate with and contaminate the shale formations.1 Some plays in Western Canada have low CO2 but enough H2S to require treating. Thus, after removing the NGLs, there are many situations in which the shale gas may still need to be treated to pipeline specifications, at least for sulfur content.

Difficulties posed by shale gases. The challenge in treating such gases is the very low H2S-to-CO2 ratio and the desire to meet, but not exceed, pipeline specifications on CO2 content. In terms of cost and effectiveness, the solvent of choice for H 2S removal and CO 2 slip is N-methyldiethanolamine (MDEA) used in a traditional gas treating plant. But how does one go about taking the H2S content from, for example, 100 ppmv down to 4 ppm without taking out excessive CO2 at the same time? Another related issue is what to do with the acid gas from the amine unit, since it will likely be of substandard quality for a Claus plant. This article uses specific examples to show, quantitatively, how various process plant parameters affect selectivity and, in particular, the ability to treat a variety of shale gases to pipeline specifications. Solvent selection, strength, temperature and circulation rate, as well as the type and quantity of internals used in the contactor, are some of the process parameters and design variables considered. Problem-solving with trays. A new tactic is to use multi-

pass trays even when, hydraulically, a single-pass tray is more than adequate to handle the flows. The key is to understand

that trays operating in the froth vs. spray regimes have radically different mass-transfer performance characteristics. A critical element in the underlying analysis is the availability of a realmass and heat-transfer rate-based simulation capability, because the selectivity issue is intimately tied to the separation taking place from a mass-transfer rate perspective. Ideal stages are incapable of dealing with this properly because, no matter how embellished by efficiencies and residence times, an ideal or equilibrium stage is completely oblivious to the effect of hydraulics on mass transfer. Rather than devoting column space to discussing what a masstransfer rate model is and how it works, we will instead present a set of case studies and simply refer interested readers to a previous Hydrocarbon Processing article2 for model details. However, it will be important in what follows to understand this fact: H2S absorption is a process controlled by resistance to mass transfer in the gas phase, whereas CO2 absorption is liquid-phase-resistance controlled. Therefore, whatever can be done to lower gas-phase resistance and increase liquid-phase resistance will improve H2S pickup and increase CO2 slip. The reaction between CO2 and MDEA is so slow that reaction kinetics play a very minor role in determining CO2 absorption rates. Carbon dioxide and hydrogen sulfide absorption are controlled strictly by the mass-transfer characteristics of the specific trays or packing under the hydraulic conditions being used. TRAYS OPERATING AT LOW LIQUID RATES

During 2007 and 2008, several plants were found to be producing gases with unbelievably low concentrations of H2S and astonishingly high CO2 slip values. These values were far outside the range suggested by any simulator, whether mass transfer ratebased or ideal-stage. In each case, the absorber contained trays. More importantly, the weir liquid load (volumetric flow rate of solvent per unit length of weir) was always quite small.3 Later, performance data was found for six more plants also operating at low weir liquid loads and, as Fig. 1 shows, the data from all nine plants show remarkable quantitative consistency with, and support for, the spray-regime explanation.4 Froth vs. spray regime. The experimental data from which

the fundamental mass-transfer coefficient correlations are drawn in a mass transfer rate-based model all corresponded to trays operating in the froth regime, in which the biphase on the trays is a frothy liquid containing a dispersed gas. However, the trays HYDROCARBON PROCESSING JANUARY 2012

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in these nine low-weir-load instances were all operating in the spray regime, with some operating with essentially pure sprays (left side of Fig. 1) and others with mostly froths but with a modicum of spray (right side). In froths, the liquid is continuous and the gas is dispersed as large and small gas bubbles and jets; in sprays, the liquid is dispersed as droplets (about 1 mm in diameter in aqueous systems) bouncing across the tray and finding their way into the downcomer through a continuous gas phase. Hydraulically, the flows are radically different, and so is the mass transfer. The spray regime has much higher liquid-phase resistance (to mass transfer) because, internally, the liquid drops are almost stagnant. The lack of mixing produces lower CO2 absorption rates (remember: CO2 absorption is liquid-phase controlled)—i.e., increased CO2 slip. On the other hand, gas-side mass transfer is much improved because of the highly turbulent flow of gas around the drops— hence, better H2S absorption. This scenario is completely consistent with what was seen in the performances of all nine plants. Thus, to improve H2S removal and to slip more CO2, trays should be operated in the spray regime if possible. Spray regime challenges. It is unfortunate that, in the

distillation community, tray operation in the spray regime has a bad rap. However, the poor reputation is the fault of many tray designers who repeatedly fail to use a seal pan to ensure that downcomer bottoms are positively, hydrostatically sealed (vs. dynamically sealed).


Alberta North Texas

British Columbia

North Texas India North Texas


North Texas

The correction factor for spray-regime operation varies with weir liquid load.


FIG. 2


1.7 1.6 1.5 1.4 1.3


16 18 20 Number of absorber trays


1.2 24

Adding trays does not meet the treating specification with MDEA at design rates.

I JANUARY 2012 HydrocarbonProcessing.com


12 CO2 in treated gas, mol%


H2S leak, ppmv

20 19 18 17 16 15 14 13 12 11 10 12

The gas plant in question is one of three units in Texas between Dallas and Houston intended to process gas from fields in the Barnett shale. As built, this particular plant was intended to treat 330 million standard cubic feet (MMscfd) of gas containing 750 ppm H2S and 2.5% CO2 at 960 pounds per square inch absolute (psia) to pipeline quality—i.e., 4 ppmv H2S and < 2% CO2. The absorber was designed with 12 single-pass valve trays using an equilibrium-stage simulator and assumed tray efficiencies. From startup in mid-2009, the plant has consistently failed to produce on-specification gas at more than 60% of the nameplate production capacity, even with reboiler and circulation pumps running at full capacity. The generic MDEA solvent was gradually spiked with a stripping promoter, allowing it to treat 240 MMscfd, or 73% of capacity. However, the internals were inadequate to move beyond this limit, and a revamp of the tower—perhaps even a new and taller column—was required. Literally hundreds of cases were run using a mass transfer ratebased amine simulator to determine the right course of action. Focusing on the absorber, the tray count was varied from 12 to 26, and solvent rates, amine strength, gas temperature and solvent temperature were varied. Consideration was given to tray type and design, the use of structured packing and even a combination of packing and trays in the same column to achieve the nameplate rate with on-specification gas. The results were somewhat surprising and very educational. Simulation results. Traditional thinking would suggest that, if a plant is not meeting treating specification, a higher solvent circulation rate and a more aggressively reboiled regenerator should improve treating. However, in the present case, the oil

Weir load FIG. 1





8 6


4 2 0 12

FIG. 3


16 18 20 Number of absorber trays


CO2 in treated gas, mol%

Spray-regime correction factor

New Mexico

A good seal prevents gas from blowing up the downcomers (rather than through the tray deck), causing massive entrainment of liquid and an undeserved bad reputation. In fact, in the spray regime, trays having positively sealed downcomers actually have higher jet flood capacity than conventionally operated trays. Attempting to seal downcomers dynamically, at very low liquid rates, is an invitation to failure. The gas-treating benefit of the spray regime was the subject of a 1981 patented tray design.5, 6 However, the patented design failed to gain popularity, perhaps because of the limited area of application 30 years ago. Nonetheless, the spray-regime operation of trays has promising application in shale gas treating today.

H2S leak, ppmv


1.4 24

The use of a stripping promoter meets treat but removes too much CO2.

LNG/GAS PROCESSING DEVELOPMENTS flow to the regenerator reboiler and the circulation rate through the unit were already at equipment limits. Furthermore, the solvent was already at 50 wt% MDEA and contained a stripping promoter, so only a small increase was possible by raising the MDEA strength by 5 wt% or 10 wt%—certainly not enough to increase performance significantly. One of the most influential parameters was simulated to be the raw gas temperature. (Solvent temperature had a much smaller effect because the gas-to-liquid ratio was high in this plant.) However, significantly reducing the gas temperature would have required a large gas heat exchanger, and any achievable lower temperature was found to be insufficient to allow treating at the design rate. Thus, the focus shifted to the tower internals. Using the right tray design. The preference was to use

generic MDEA rather than a specialty amine. Fig. 2 shows that, with generic MDEA, adding trays will indeed lower the H2S leak into the treated gas, but not nearly enough to meet the H2S specification. Note that the weir load in this case is 65 gallons per minute per foot (gpm/ft), requiring application of a small correction for a small amount of spray. However, the problem with the absorber is that, the more trays there are, the more CO2 is removed. Already twice as much CO2 as necessary is being removed from the gas. Solvent capacity is being used to remove the wrong component (CO2) instead of the noncompliant component (H2S). No matter how many trays are used in this absorber, generic MDEA will not allow the gas specification to be met at design rates.


■ Contrary to urban legend, entrainment

rates and tray capacity do not have to be negatively affected by the sprays that accompany low weir liquid loads. Using a stripping additive would permit the originally intended gas rate to be processed to pipeline specifications, as Fig. 3 shows. However, mass transfer rate-based simulation shows that at least 20 absorber trays would be needed, and even if 20 trays could be shoehorned into the existing shell, twice the necessary amount of CO2 would be removed. It turns out that a moderate crimp structured packing could be used effectively in this particular column, achieving less than 1 ppmv H2S and 1.95% CO2 in a 35-ft bed, but only with an amine solvent containing a stripping promoter. With generic MDEA, simulation showed that 6–7 ppmv H2S was the best that could be achieved, albeit with 1.9–2.0% CO2. Hydraulically speaking, one-pass trays are perfectly adequate for handling the gas and liquid flows in the absorber. However, if two-pass trays were installed, the 65-gpm/ft weir load would drop to about 40 gpm/ft, and a significant benefit to both H2S removal and CO2 slip would result. Furthermore, rich-solution loadings are quite modest, so the solvent has more capacity than is being used. This situation suggests that, if the solvent rate were reduced to below the plant limit, even lower weir load and better







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I 47



7 H2S CO2

H2S, ppmv


2.05 2.00


1.95 4 1.90 3 2 200

FIG. 4

1.85 250 300 Circulation rate, gpm

CO2 in treated gas, mol%


1.80 350

Processing 330 MMscfd at reduced circulation rate using 20 two-pass trays with 50 wt% generic MDEA.

H2S removal and CO2 slip would result. Fig. 4 shows simulated treating results for a 20-tray absorber containing two-pass trays as a function of solvent rate. This absorber is simulated to handle the full-design gas flow (330 MMscfd) using only generic MDEA at just 70% of the nameplate solvent rate. The keys are using mass transfer ratebased simulation, and knowing how tower internal details (e.g., tray passes) affect the absorption process. This kind of technical sophistication allows a simulation model to be converted into a “virtual plant.” An absorber that was completely unable to meet design criteria—no matter how many trays or how much packing it contains—has been transformed into a success. As a backup plan, the danger of a small margin for error in meeting the H2S specification can be mitigated by using a specialty amine to achieve < 0.5 ppmv quite easily. Shale gas can be very challenging to treat. However, mass transfer rate-based simulation and appropriately specified and designed tower internals can make shale gas treating no harder than treating any other gas. Without both ingredients, though, treating shale gas can be a guessing game. TREATING SHALE GAS FROM BRITISH COLUMBIA

This particular example has 26 ppmv H2S and about 1.1% CO2, so the gas needs to be treated for H2S while allowing as much CO2 slip as possible, since CO2 is already below pipeline specifications. Due to the very small amount of acid gas needing to be removed, the absorber has only 12 trays, and plant data indicate that the rich solvent is lightly loaded. The treated gas is below 4 ppmv (no measurement is available), and the unit is slipping about 80% of the CO2. The weir load is about 30 gpm/ ft, so the amount of spray is a significant fraction of the total biphase on the tray. Simulation with no adjustment for low-weirload (spray-regime) operation suggests a CO2 slip of 54% with 1 ppmv of H2S slip. However, when proper account is taken of the hydraulic operating region in which the trays are operating (spray regime), the simulated CO2 slip is 78% vs. 80% measured. The H2S treat is 1.3 ppmv, well below the 4-ppmv specification. Obviously, the tray hydraulic operating region has a profound effect on treating. In particular, selectivity is a very strong function of a tray’s hydraulic operating region. The simulations are truly outof-the-box predictions because no input was used beyond tray construction details and basic plant flows. Nothing was tweaked to force a match to performance data. 48

I JANUARY 2012 HydrocarbonProcessing.com

With the wrong modeling tools, shale gas treating units can be very challenging to simulate and, therefore, challenging to build with any reasonable assurance of performance. The difficulty lies in the very low H2S content of shale gases, which leads to low liquid-to-gas flowrate ratios in amine contactors. A critical and essential element in reliable tower design for shale gas treating is a solid mass transfer rate-based simulator, because tray hydraulics profoundly affect not just pressure drop; they also impact mass transfer and the very separation process itself. Ideal stage calculations are oblivious to what is actually in the column, let alone the mode of operation. Under conditions that are common in shale gas treating, trays will often have to be operated in the spray regime, where care must be taken on the part of tray designers and design engineers to ensure that downcomers remain positively sealed against massive bypassing of gas. However, even when trays operate with froths, there is great potential advantage to be gained from contriving methods to force operation into the spray region, and the more spray-like the biphase, the greater the potential advantage in terms of enhanced selectivity. Contrary to urban legend, entrainment rates and tray capacity do not have to be negatively affected by the sprays that accompany low weir liquid loads. However, tray designers must be attentive to the need for positive downcomer seals, preferably through the use of recessed seal pans beneath the downcomers. Multi-pass trays are an under-appreciated but powerful weapon that can be brought to bear in amine unit design to meet the unique treating challenges offered by shale gases and other gases requiring small liquid flows to treat large volumes. HP LITERATURE CITED Hunter, J. C., “The New Albany Shale from an Antrim Shale Operator’s Perspective,” RPSEA/GTI Gas Shale Forum, Des Plains, Illinois, June 4, 2009. 2 Weiland, R. H. and N. A. Hatcher, “What are the benefits from mass transfer rate-based simulation?” Hydrocarbon Processing, July 2011. 3 Weiland, R. H., “Tray Operating Regimes and Selectivity,” Laurence Reid Gas Conditioning Conference, Norman, Oklahoma, February 22–25, 2009. 4 Weiland, R. H., N. A. Hatcher and J. L. Nava, “Designing Trays for Selective Treating,” SOGAT 2010, Abu Dhabi, UAE, March 28–31, 2010. 5 Resetarits, M., Personal communication, 2008. 6 Sigmund, P. W. and K. F. Butwell, US Patent 4,278,621, July 14, 1981. 1

Ralph Weiland founded Optimized Gas Treating Inc. in 1992 and has been active in Canada, Australia and the US in basic and applied research in gas treating since 1965. He developed the first mass transfer rate-based model for amine columns for Dow Chemical and is responsible for the development of the Windows-based ProTreat process simulation package. Dr. Weiland also spent 10 years in tray research and development with Koch-Glitsch LP, Dallas, Texas. He earned BASc and MASc degrees and a PhD degree in chemical engineering from the University of Toronto.

Nate Hatcher joined Optimized Gas Treating Inc. as vice president of Technology Development in 2009. He is responsible for making improvements and adding functionality to the ProTreat gas treating process simulator. Mr. Hatcher has spent most of his 16-year career involved with sour-gas treating and sulfur recovery, first in design and startup and later in plant troubleshooting, technical support and process simulation development. He is a member of the Amine Best Practices Group and serves on the Laurance Reid Gas Conditioning Conference advisory board. Mr. Hatcher received a BS degree in chemical engineering from the University of Kansas and is a registered professional engineer in the state of Kansas.



Innovative APC boosts LNG train production APC application yields significant operability, economic benefits A. TAYLOR, Apex Optimisation, Adelaide, Australia; and S. JAMALUDIN, Woodside, Karratha, Australia


he appropriate use of advanced process control (APC)— specifically, multivariable predictive control (MPC)—has been well established in the hydrocarbon processing industry over multiple decades, and it is widely considered an essential contributor to production maximization on liquefied natural gas (LNG) trains. If correctly applied, APC software delivers more efficient operation of existing hardware assets and essentially provides a “cruise control” for the control room operator. The Woodside-operated Karratha Gas Plant (KGP) has been progressive in the application of APC across all major process units, generating sustained benefits. Although the site is a mature APC user, there is a continual focus on innovation and design evolution to further improve APC benefits. This article describes the implementation of APC on an LNG liquefaction train. Several generic APC project aspects are investigated, such as the use of a dynamic simulator and automated step testing to aid development. Also, details of the project’s significant operability and economic benefits—including a 4,000% return on investment—are discussed with commentary on whether this success has been sustained beyond the “honeymoon” period.


Woodside engaged Apex Optimisation to assist with a revamp of the existing APC on LNG train 4 (LNG4) and the implementation of a new APC on LNG train 5 (LNG5). The project was a collaborative effort, with both parties heavily involved in the design, implementation, commissioning and post-audit of the new APCs. The implementation kicked off in March 2010 after a functional design specification phase. The revamped LNG4 APC and the new LNG5 APC were commissioned in May 2010 and September 2010, respectively. A successful site acceptance test signaled handover to site support engineers in October 2010. Challenges to development. The execution of the project

was challenging due to a range of factors: • The design evolution significantly pushed the previous project’s boundaries. Additional compressor power-management handles were included, the site electrical power-generation spinning reserve and fuel gas system capacity limits were added (these global constraints are relevant to both trains), and a more sophisticated approach to optimizer functionality was adopted. Hence, the scope of the modeling and custom functionality required was substantially different from that of the previous LNG4 APC application.

• The new applications are relatively large, with each having over 20 manipulated variables (MVs) managing more than 60 controlled variables (CVs) and some complex interactions (i.e., relatively high model density). • Parts of the process are highly nonlinear in their behavior, and this can limit the applicability of linear APC technologies. Improved performance was needed during lower production conditions (e.g., turndown or hot summer temperatures), and this required some innovative use of transforms, gain scheduling and automatic logic to manage variable usage. Dynamic simulation was leveraged to develop the gain scheduling relationships. • As the existing LNG4 APC had been unused for over a year, there was limited operator expertise with APC on the LNG4/ LNG5 distributed control system (DCS) panel. This situation required careful management of the reintroduction of APC and operator training. • The LNG5 train was relatively young, with a limited operating history. Furthermore, its operation was very different from that of LNG4, despite the equipment design being essentially identical. Mechanical changes to the LNG5 train during the execution phase of the APC project significantly changed the train operation and reset the LNG5 APC design needs. The project engineers had to remain flexible to adapt to the changing basis while maintaining the project schedule. • Interfacing to some of the compressor packages required an exotic approach. In particular, one key compressor handle was hosted on a separate DCS network on the other side of the control room. This context required careful software design and operator training to ensure that the final mechanism was robust and intuitive to both DCS operators. • Automatic step testing was adopted in order to reduce the duration of the step-testing phase; this had not been previously attempted onsite. • An aggressive schedule was required to commission two large applications within seven months, which kept the intensity high throughout the duration of the project. Fig. 1 shows a schematic illustrating the process design for the two liquefaction trains. These challenges were overcome through teamwork among the participants. Close operator involvement was critical to project success, as this fostered ownership of the project and ensured that each process control improvement implemented was intuitive for the operators and appropriate for the widest range of process conditions. HYDROCARBON PROCESSING JANUARY 2012

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One of the major APC benefits delivered is improved consistency in how the process is managed. To realize this benefit via sustained APC usage, consistency in how the APC is operated is paramount. Therefore, thorough operator training is essential to the project process. Fig. 2 shows Woodside DCS operators at work in the control room. USE OF DYNAMIC SIMULATORS TO ASSIST MODEL DEVELOPMENT

In recent years, the use of a dynamic simulator (i.e., an operator training simulator, or OTS) has been promoted by advocates as a more efficient way of developing APC. The ability to speed up real time, avoid real-life plant reliability and load disturbance impacts, reduce engineering support requirements, and potentially complete the APC development well before the plant is commissioned makes the OTS very appealing to cost- and schedule-focused customers. These factors prompted Woodside to investigate the use of an existing OTS to assist with the conceptual design and initial (“seed”) model for the automated step-test phase. While the OTS is typically fit for the purpose of investigating an APC optimization strategy and controller structure, is it appropriate for APC model development? One can build an OTS to varying levels of fidelity (with cost implications), and the main objectives are typically:1 • Enabling thorough DCS and emergency shutdown system checkout and verification before construction Rich gas HMR


FG to PC comp GTs







Process design for LNG4 and LNG5.

N2 rejection



Scrub column


FG LNG to storage

• Providing useful operator training on the process with the target system interface • Providing a useful “what if?” tool for engineering analysis of process changes. Ascertaining OTS fidelity. To achieve these objectives, the

OTS requires a level of fidelity that is well practiced and accepted by OTS developers. However, a standard OTS may not have the fidelity required for complete APC model development; what is required is a function of both the APC modeling needs (the APC design) and the nature of the process included in the APC scope. Even if it is identified as an OTS objective up front, the distant APC topic may struggle to justify a costly increase in the OTS fidelity among more traditional construction project needs. The question then becomes, “How can it be known if the OTS has the required fidelity?” This question is not an easy one to answer unless an operating plant can be used as a datum, or unless the process is extremely well understood from a modeling perspective and the required fidelity exists. In our LNG liquefaction APC example, the OTS system was developed alongside the construction project, with traditional objectives in mind and well before APC was considered. The development of the OTS was given heavy focus (including post-commissioning improvements to OTS accuracy in selected areas), with high acceptance of the simulator’s value. When using the OTS for the APC model development, we found that the thermodynamics-related models were reasonably accurate at base-case production rates. However, there were discrepancies around many of the ΔP-related models (especially those associated with complex devices such as hydraulic turbines with multiple flow elements) and turndown-related models (such as those associated with flow regime changes experienced inside the spiral-wound cryogenic heat exchanger). Given the exotic nature of the cases where accuracy was lacking and the relative importance of these items to the traditional OTS objectives, this is not a surprising outcome from a traditional OTS used outside of its original purpose. The value of the OTS in our LNG APC case was essentially limited to the actions listed below: • Formulating the optimization strategy and controller structure • Being able to interrogate turndown cases, which are relevant for hot-weather operation, without suffering production losses on the plant or needing to contemplate a second step test in more difficult summer conditions—thus, providing valuable data on relative gain changes, which was used in the gain scheduling logic • Providing useful, initial models for the automatic stepping tool. As the new APC design was different in both DCS control basis and scope, the previous model could not meet this need in all areas. Benefits of simulation. A dynamic simulator of typical

FIG. 2


Woodside DCS operators at work.

I JANUARY 2012 HydrocarbonProcessing.com

fidelity (OTS or desktop engineering tool) can be useful in verifying an APC design concept in terms of control and optimization strategies. This need is more relevant for complex processes where the pre-APC operation does not exploit all the available degrees of freedom and some methodology needs to be developed. The APC model accuracy required for accurate model development and full APC benefits would be much higher than that required for strategy verification. A complete OTS-based APC model was developed as part of the functional design phase to support the automated step test. After the final model was verified post-commissioning, a comparison was performed to assess the accuracy of the OTS-based model. The results in several key areas are presented in Table 1.

LNG/GAS PROCESSING DEVELOPMENTS In summary, the knowledge gained from using an OTS for APC model developments (as distinct to APC design and optimization strategy) reinforces the following guidelines: • Understand the relevant accuracy of the OTS well. There are obvious implications for developing APC on young or difficult OTS processes prior to plant commissioning. In some instances, the OTS has relevant accuracy inherently (e.g., the C3 splitter example, where the distillation models are the key aspect2). In other areas, the important APC needs are not necessarily aligned with key OTS objectives. • Understand the value of using the OTS in APC development; i.e., is it prohibitive to step test on the real plant for operational or economic reasons? • Do not underestimate the value of working on the real plant and interacting with operators for developing an operations understanding (as distinct to a process understanding) and cultivating APC understanding. • Always be prepared for some model error when commissioning the APC on the real plant, and allocate sufficient time to resolve any problems. USE OF AUTOMATED STEP-TEST TECHNIQUES

Automated step-test techniques have been promoted in recent years as a way of providing a rich data set in a short period of time, thereby reducing project cost. Also, simultaneous testing of multiple MVs could improve the accuracy of the gain ratios that are important to the performance of the application. This LNG liquefaction APC project was the first incidence in which the site had used this technique as the primary step-test approach, after successful testing on the liquid petroleum gas (LPG) fractionation unit suggested it would be a time-saving option. Despite the best endeavors of the project team, the LNG train experience was somewhat different, with the net result being neutral relative to a traditional, manual step test. The reality was that this particular LNG liquefaction process was not well-suited to this technique, for the two reasons listed below: 1. The daily variation due to ambient temperature swings is six times the maximum MV step size allowed for the test. The automated tool works purely on process feedback, whereas anyone operating the plant knows what moves have to be made before the sun comes up. The manual test is superior in this case, as the tester can plan moves using all information available, not just APC variables. Thus, when using the tool as intended, the moves required to control the process swamped the random steps required for model identification. 2. Also, the extent of the load disturbances encountered during a normal day demands both the need for minimal optimizer action and the inclusion of extra steps in addition to the automated steps. For other processes where this is not the case, and manual steptest costs are greater, this approach may offer a tangible reduction in the step-test duration.

test. In our LNG example, the superseded DCS controls provided a high level of optimization that had to be matched during the step test. The automated test must accommodate this need with some sort of mild optimization. • It may be useful to automatically change step direction if a full step size is not feasible due to potential limit violations. If partial moves are applied, additional steps may be required to achieve the same data quality. • As there can be a need to make extra moves on a real plant, it may be desirable to include all moves made during the step test— not just those made by the automated tool—in the model identification approach, as a means of reducing the total test duration. • Real-time model identification can be very useful, but one should not rely only on automated model identification to signal that testing is complete. In one instance, this approach produced some false negatives, which would have prolonged the test further if additional identification was not undertaken using traditional approaches—i.e., manual data grooming, careful slicing, and finite impulse response (FIR) generation over multiple times to steady state (TSS). • Engineers should not be required to work more intensely than a manual step test in order to manage the automated testing. Keeping in mind that the traditional approach offers some additional value: 1. Time for detailed discussions with operators at the panel is very effective from both a “public relations” and training perspective. 2. Time to observe the plant behavior and “experience the challenge for the APC” provides useful insight into how the APC should act and sets helpful expectations for the model identification. Unfortunately, this valuable experience is generally negated by automated testing tools, which step multiple MVs simultaneously as the CV responses can no longer be seen by the eye. 3. Time to consider DCS control servo response and make repairs early can greatly improve the final result. It is widely regarded that most efficiency tools added to a wellproven methodology are no replacement for sound engineering judgment. Generalizations about efficiency improvements will be tested by the more challenging APC projects. One needs to have confidence in significant efficiency gains to warrant deviation from the trusted methodology, especially when the payback on these projects is already substantial. CUSTOMIZATION OF APC APPLICATIONS

Woodside has nearly 15 years of experience with APC applications in the relatively demanding environment of an integrated TABLE 1. Accuracy assessment results for final OTS-based APC model

Test automation results. Based on our cumulative experi-

ence with a range of automated step-test techniques, our conclusions from test automation are set out below: • Using the available APC model as a true model identification “seed” model (as opposed to simply a model used by the APC to manage the process during the test) may considerably speed up the model development process. A further enhancement would be the ability to assign confidence to sub-models to assist the initial model identification. • With some processes, it is not viable to switch off the optimizer action for long periods, much less for the duration of a step


Plant area

Proportion of final models closely resembling OTSbased models


Scrub column


Good form and gain from OTS models



OTS gains were regularly less than half of plant test gains (10% at times)

Hydraulic turbines


OTS gains were sometimes inaccurate by two orders of magnitude


I 51



production facility. The context is demanding in the sense that personnel turnover is high at the remote site, and the costs of poor performance are severe. Accordingly, effort is required to maintain appropriate skill levels at the site. This experience has proven the value of appropriate APC customization to improve availability and robustness. Indeed, the inability of the previous APC application to accommodate the full range of operations was one of the main reasons for its demise. A few examples of how the generic APC software was augmented are discussed below. Gain scheduling for turndown. Analysis of previous

APC performance and OTS scenarios confirmed significant gain changes at reduced production rates. These changes demanded custom logic to manage gain scheduling, according to production rate ranges using discrete gain multipliers. (Continuous formula-based gain scheduling was not preferred due to the risk of producing ill-conditioned matrices.) The logic also provided some automatic shedding of specific MVs and CVs during turndown to accommodate the unique operating context. Model adaptation for hydraulic turbines. The power extraction from the hydraulic turbines is akin to climbing to the summit of a hill, with constraints applying a ceiling on how high one can climb. The model gains are very much a function of the status of the surrounding DCS controls, and if the alternative flow path opens up (the Joule Thompson [JT] valve), the wicket gate is moved in the opposite direction to maximize power extrac-

tion (i.e., one is on the opposite side of the hill and needs to walk in the other direction to climb it). In the past, this scenario had constituted a challenge for the APC that was avoided by instructing the operators to ensure that the JT valve was shut before giving the wicket gate control to the APC. However, it was still possible to suffer load disturbances, which bounced the process onto the opposite side of the gain inflexion point. The results were not positive. With the addition of simple logic to flip the gain sign and drop/ activate specific constraints, the new APC has improved robustness by allowing the operators to give the hydraulic turbine control to the APC, regardless of the DCS control state. The APC will honor the correct constraints with appropriate wicket gate moves, and will walk the process over to the “correct side of the hill” when feasible. Product price-driven optimization. Another feature of

the new APC design is the ability to specify product prices and use them to dictate the subtleties of the optimization toward either maximizing LNG production or LPG extraction. This arrangement is different from simply specifying maximum LNG or maximum LPG, as each of the relevant MVs has differing effects on the yield of each product. It is useful to provide some “shades of gray” in terms of the optimization options. Aside from a purely economics-driven optimization, the APC has maximum LNG and maximum LPG modes to assist logistics needs without sacrificing valuable production (e.g., tank-top scenarios that affect only one product). PROJECT RESULTS

Refrigerant compressor power balance, pre-APC 25 MR helper power C3 helper power


20 15 10 5 0 3/24/09 FIG. 3



4/8/09 Date



Power consumption of the primary compressors pre-APC.

The overall results of the project were exceptional, given the challenges faced. Results included: • Excellent operator acceptance of all the developments implemented during the project (i.e., DCS control improvements, instrument repairs and APC commissioning), with APC uptimes consistently greater than 97%. Operator feedback shows that the new APC makes objectives easier to achieve. • A tangible contribution to improved reliability as a result of the APC maintaining the process within constraints on a minuteby-minute basis. In particular, the APC manages some difficult operating envelope constraints associated with the large axial compressors employed in the liquefaction process. Prior to the APC, manual management of this relatively tight feasible space, coupled with the production changes driven by diurnal swings, left the DCS operators under continual pressure. Frequency plot of % of technical maximum capacity, pre- and post-APC

Refrigerant compressor power balance, post-APC


20 15 10

0 4/13/11 4/15/11 4/17/11 4/19/11 4/21/11 4/23/11 4/25/11 4/27/11 Date


7 6

With APC, one year post-commissioning Without APC With APC, post-commissioning

5 4 3 2


FIG. 4


MR helper power C3 helper power

Frequency, %


Power consumption of the primary compressors post-APC commissioning.

I JANUARY 2012 HydrocarbonProcessing.com

1 0 70 FIG. 5


80 85 90 95 Technical maximum capacity, %



Comparison of production vs. technical maximum capacity.

LNG/GAS PROCESSING DEVELOPMENTS • The production increase achieved with the same process equipment represents a decrease in specific energy consumption and a relative reduction in carbon footprint for this important clean energy-producing process. • The project was completed on schedule and within budget, despite an evolving design datum being prevalent throughout the execution. • The APC benefits delivered a significant boost to the bottom line for North West Shelf Joint Venture Partners, with a 3%–5% increase in LNG4/LNG5 production (depending upon ambient conditions) and a 4.7% increase in LPG production verified. This production increase delivered an overall project payback of less than two weeks, or a return on investment of 4,000%. • At the 2011 Process And Control Engineering (PACE) Zenith Awards, the project won the Oil & Gas category and the Project of the Year Award ahead of 50 competing projects. The LNG production benefits are best illustrated by the reduction in compressor power giveaway, which is an inherent characteristic of the process design. That is, production is either limited by the helper motor power on the mixed refrigerant (MR) compressor or the propane (C3) compressor. The amount of spare compressor power not applied to the process represents a production loss. Fig. 3 shows power consumption of the primary compressors before the APC. Following the commissioning of the new APC, the higher average power consumption was a significant contributor to the increased production capacity. Fig. 4 shows power consumption of the primary compressors after APC commissioning.


It is important to note that the project benefits have been sustained one year later, with no deterioration in performance or in operator satisfaction detected. Fig. 5 shows a comparison of production vs. technical maximum capacity. This project demonstrates how the appropriate use of APC technology can provide a tangible and sustained improvement in plant profitability and operability in a cost-effective manner. HP LITERATURE CITED Stephenson, G. and L. Wang, “Dynamic simulation of liquefied natural gas processes,” Hydrocarbon Processing, July 2010. 2 Alsop, N. and J. M. Ferrer, “Avoiding plant tests with dynamic simulation,” Hydrocarbon Processing, June 2008. 1

Andrew Taylor is a principal consultant with Apex Optimisation, based in Australia. His responsibilities include all aspects of APC application design, implementation and maintenance. In his 20 years of experience, he has contributed to over 100 APC applications. Previously, he was employed as a consultant with Honeywell in South Africa and the UK and with Mobil in Australia. Mr. Taylor holds a BE degree in engineering science from the University of Auckland and is a chartered professional member of Engineers Australia.

Saifullah Jamaludin is a senior process control engineer at Woodside Energy Ltd. and has 12 years of experience in the LNG industry. He was previously employed by Petronas in Malaysia. Mr. Jamaludin has published numerous papers for technical journals and international industrial conferences. He contributed to the development of the first LNG train automatic cool-down advanced controller, and has led the design and implementation of multiple APC applications. He holds a BS degree in chemical engineering from the University of Edinburgh.

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Apply new enhanced tubes to optimize heat transfer in LNG trains New developments for heat exchangers reduce capital and plot size of key equipment B. PLOIX, Technip France, Paris, France; and T. LANG, Wieland-Werke AG, Ulm, Germany


Tr 4


Ni ge

LNG plant capacity, million tpy

pplications for extruded lowfinned (LF) and externally and internally enhanced tubes are widespread in multiple industries, ranging from the air-conditioning and refrigeration, heating, automotive and power industries, as well as the hydrocarbon processing industry (HPI). A few selected examples are shown here: • Enhanced boiling and condensation tubes for packaged chillers • Inner grooved tubes for coil heat exchangers in the air-conditioning and refrigeration industry • Enhanced tubes for the hydrocarbon processing industry • Enhanced tubes for power steering oil FIG. 1 Enhanced internal and external LF tubes. cooling in the automotive industry. For the HPI, distinct enhanced tubes for boiling and condensing, as well as sin9 gle-phase heat transfer services have been derived from standard Qatar mega 1 8 LF tubes (Fig. 1). The enhanced tubes are typically manufactrains 7 tured from plain tube-based material by an extrusion cold-rolling 6 process. The fins on the outside, and on the inside are integrally connected to the tube wall. 5 r1 A wide range of proven references for both LF and enhanced 4 Yemen aT eri g i N tubes exists in the HPI from decades of refining, petrochemi3 cal, chemical and gas processing applications.2–4 Standard tube 2 Skikda materials are copper-nickel carbon and low-temperature carbon 1 Camel steels. Now solutions are also available in low-alloy carbon steel, 0 1960 1965 1970 1975 1980 1985 1990 1995 2000 2005 2010 2015 as well as stainless steel (SS) and titanium (Ti). These technoloSource: Technip gies resolve both capacity and plot-space limitations for existing plants and provide compact or most efficient solutions for new FIG. 2 Evolution of LNG plant capacity (Technip references). plant design. Very attractive applications are identified in liquefied natural gas (LNG) and ethylene plant due to the drastic recent increased plant capacity (Figs. 2 and 3). two-phase flow conditions, the overall heat transfer benefit leads to substantial overall benefits. At the same time, the tube-side ENHANCED HEAT TRANSFER TECHNOLOGIES pressure drop increases. However, it typically does not exceed the The thermal advantage of the externally and internally heat transfer improvement, allowing for enhanced shell-and-tube enhanced tubes vs. a plain tube, in a shell-side propane boiling heat exchangers within allowable pressure drop limits. application is between a factor 2 and 3 of the boiling heat transfer The key benefit of the externally and internally enhanced tube coefficient. Together with an improved tube-side performance is the capability of a superior operation for an external boiling ranging between a factor 1.6 to 2.4, depending on single- and application at low-temperature approaches—down to 2°C and HYDROCARBON PROCESSING JANUARY 2012

I 55



exchangers. Special attention is given both to the thermal and mechanical design, along with an improved kettle design, such as:7 • Proper fluid distribution at inlet and outlet • Verification of liquid entrainment especially for suction line to compressor State of the art thermal design tools from Heat Transfer Research Inc. (HTRI), as well as other advanced heat transfer are used as convenient thermal design tools for enhanced heat transfer solutions.8 In a similar way to the enhancement of reboilers, enhanced heat transfer solutions have been developed for horizontal con1,600 Yansab KSA 10th complex densers, with shell-side condensation of pure streams and tube1,400 Iran side cooling water. A typical solution is with an enhanced tube 1,200 having an external LF structure in combination with an internal Ras Laffan helical fin structure. 1,000 Qatar Both enhanced heat transfer technologies have been made 800 available and qualified through local testing for horizontal shell600 and-tube type reboilers and condensers. During two joint industry and academia research projects, JOULE III and AHEAD, 400 funded by the EU, fundamental research and qualification have 200 been conducted.5,6 Key activities have been the development 0 and characterization of enhanced shell-side nucleate boiling 60s 70s 80s 90s 2000 2010 Years structures, especially at low-temperature approaches, as well as tube-side enhancement structures for both single-phase gas and FIG. 3 Evolution of ethylene plants capacity (Technip references.) liquid and two-phase condensate heat transfer. The different steps from lab-scale testing to industrial application stretched over a period of almost Cracked gas compressor one decade. The pre-requirement for the time being in these applications is for clean Hydrogenation reactor Tail gas refrigerant and process fluids. Cold box expander 1-4 5 For base-load LNG plants, the enhanced Ethylene Furnaces heat exchanger technologies are highly attractive within the propane pre-cooling Feed C2 cycle. The major application for enhanced splitter Chillers boiling tubes is for the main propane refrigDryer erant chilling train with cooling/condensaTLEs tion of natural gas (NG) or mixed refrigerant (MR) on the tube side and propane Recycle ethane + C refrigerant boiling on the shell side. The 3 Quench Caustic Demethanizer tower tower Deethanizer primary application of the enhanced condensation tube is for the propane refrigerant Ethylene back-end hydrogenation process scheme based on ethane feedstock. FIG. 4 condenser with shell-side propane refrigerant condensing and tube-side cooling water. For ethylene plants, enhanced heat Air or water cooled exchanger solutions are available for the C3 ref. condenser majority of reboiler and condenser heat Propane exchangers in the cold section such as the precooling cycle C3 C2 and C3 fractionation and splitting serMain heat C3/NG exchanger chilling vices as well as the refrigerant units (as sumtrain LNG NG feed marized in Table 1). Ethylene plant capacity, thousand tpy

below where standard plain or LF tubes are no longer efficient. In various schemes and applications, the benefits can be: • Size reduction • Reduced number of heat exchangers per unit • Capacity increase or energy consumption reduction resulting from improved efficiency. A new enhanced condensing tube is available for industrial reboiler applications such as thermosiphon and kettle heat


C3/MR chilling train

Several studies have been done based on recent projects (FEED and EPC) allowing the technical and economic qualification of enhanced heat exchangers both in ethylene and LNG plants.

MR liquefaction cycle


FIG. 5


C3/MR liquefaction section.

I JANUARY 2012 HydrocarbonProcessing.com


Ethylene. In ethylene plants, there are

many heat exchangers, with very large heat transfer surface area representing around 20% of the total equipment cost of a plant.

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In parallel to the technical qualification of enhanced tubes in ethylene plants, a study was conducted to evaluate the economic interest of such solutions compared to the plain and LF tube solutions considering identical process conditions. This study concerns the C2H4 back-end (BE) hydroprocessing scheme of a typical ethylene plant, as shown in Figs. 4 and 5. The economic interest of enhanced boiling and condensation tubes is demonstrated using the key exchangers listed in Table 1. All design being conducted with maximal usage of the allowable tube-side pressure drop focus on two main goals:

Relative area per equipment

100 90 80 70 60 50 40 30 20 10 0

TABLE 1. Overview on enhanced heat transfer technologies in LNG and ethylene plants

Deethanizer C2 splitter Depropanizer C2 splitter condenser condenser condenser reboiler


FIG. 6

Low fin.

Plant type



Propane refrigerant chiller (pre-cooling cycle for NG and MR) Propane refrigerant condenser (pre-cooling cycle)


Deethanizer condenser C2 splitter condenser Depropanizer condenser C2 splitter reboiler Ethane vaporizer C2 refrigerant condenser C3 splitter condenser Propylene/propane condenser

Ethane C2 refrigerant vaporizer condenser

Enhanced condensing tube

Relative heat transfer area for items equipped with enhanced boiling tubes. 100 90 80 70 60 50 40 30 20 10 0

Relative area per equipment

120 Relative cost per equipment

• Heat transfer surface area reduction • Shell number reduction. For shell-side boiling services, Figs. 6 and 7 show the relative comparison of plain, LF and enhanced boiling tube design with the plain tube as reference. In conclusion, the average heat transfer surface area reduction is about 60%, and the average cost reduction is about 20% per equipment. For shell-side condensing services, Figs. 8 and 9 show the relative comparison of plain, LF and enhanced condensing tube design with the plain tube as reference. The average heat transfer surface area reduction is about 75%, and the equipment cost average reduction is about 65% per equipment. Additional savings come from:

100 80 60 40 20 Deethanizer C2 splitter Depropanizer C2 splitter condenser condenser condenser reboiler


Low fin.

Ethane C2 refrigerant vaporizer condenser


Relative cost for items equipped with enhanced boiling tubes.

Relative cost for items equipped with enhanced condensing tubes.

45 MW NKN, 1-pass, bundle OD = 1,500 mm, ¾-in. enhanced condensing tube, tube count: 3,745 Shellside fluid Propane ref., boiling, Tsat = 21.8°C Tubeside fluid Mixed ref., condensing, Tin/out = 1.9/-18.5°C

Relative area per equipment


Plain 19


Low fin.

C3 splitter condenser

I JANUARY 2012 HydrocarbonProcessing.com

New tube

124 79

Tube length, m Weight of shell (dry), tons

Enhanced condensing tube

Relative heat transfer area for items equipped with enhanced condensing tubes.


LF, 30 fpi 10.9

Propylene refrigerant condenser


Enhanced condensing tube

Heat duty Shell design

100 90 80 70 60 50 40 30 20 10 0

FIG. 8

Low fin.

Enhanced condensing tube

FIG. 9 FIG. 7

C3 splitter condenser

Propylene refrigerant condenser


FIG. 10

LP/MR propane refrigerant chiller. Comparison of plain, LF and enhanced boiling tube.


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LNG/GAS PROCESSING DEVELOPMENTS • Process optimization, considering the low temperature approach capabilities of externally and internally enhanced tubes and • Piping and structure reduction due to plot plan reduction, which are not included in this study.

to expand the capacity of the existing three trains from 2 million tpy (MMtpy) to 3 MMtpy per train. A new kettle-type chiller Heat duty Shell design

61 MW NKN, 1-pass, shell ID = 2,280 mm, ¾-in. enhanced condensing tube, tube count: 6,467 Shellside fluid Propane ref., boiling, Tsat = 36°C Tubeside fluid Closed cycle cooling water, Tin/out = 22.0/-31.2°C

LNG. The performance of enhanced boiling and condensing

tubes is demonstrated in two representative cases both for a propane-refrigerant chiller and condenser in comparison to standard plain and LF tubes, as shown in Figs. 10 and 11. The cases are taken from a recent LNG project. In both cases, substantial size and weight reduction can be achieved by using externally and internally enhanced tubes. Especially for the large equipment units, the benefit becomes evident when considering the whole supply chain ranging from fabrication and transportation, as well as plant aspects covering installation, operation and maintenance. A detailed techno-economic study of the two chilling trains for NG and MR showed very attractive savings in capital expense (CAPEX) and plot space, as well as capabilities for efficiency improvements or, vice versa, an attractive opportunity for capacity increase, as summarized in Table 2. Both solutions with LF and internally and externally enhanced boiling tubes have been analyzed for the two chilling trains: propane/MR chilling train and propane/NG chilling train. Each train is operating at four propane levels. For the externally and internally enhanced boiling tube, a reduction of the cold approach to 2K is feasible and considered an improved LNG plant design. Other items have been considered for the CAPEX and include heat exchanger, piping, steel structure, piping and exchanger foundation. For the standard cold approach of 3K, the externally and internally enhanced tube allows for a CAPEX reduction of 20% and 25% reduced plot space vs. a standard solution using LF tubes. Considering a reduced cold approach of 2K, the compression power is reduced by approximately 2.2% translating into approximately 1% additional LNG capacity. The additional annual income, depending on the LNG price, is far superior compared to the total cost of the chilling train. Note that the case with the enhanced tube and 2K cold approach is with 13% plot space reduction, is still more compact with the same CAPEX and is not more expensive compared to the LF case with a 3K cold approach. INDUSTRIAL APPLICATIONS


Plain 9

erence of the enhanced boiling tube dates from 2000 for a horizontal thermosiphon, C3 splitter reboiler as part of the capacity expansion of the Lyondell-Basell polypropylene plant in Knapsack, Germany (Fig. 12).8 The use of an enhanced boiling tube allowed an upgrade from 4 MW to 5 MW despite a substantial reduction of the LMTD. The cooling water return from the tubular polymerization reactor was able to be used for heating, thus avoiding the use of stream.

Propane refrigerant condenser. Comparison of plain, LF and enhanced condensing tube.

FIG. 12

Installation of horizontal thermosiphon reboiler equipped with enhanced boiling tubes in a C3 splitter of a Lyondell-Basell polypropylene plant in Knapsack, Germany.

-37°C, 6.5 bar Kettle reboiler/ condenser

29.5 °C, 20.7 bar

Compressor Ethylene Ethane

LNG plant. In 2003, the enhanced boil-

ing tube was applied for the first time in an LNG plant as part of the Qatargas debottlenecking project. The objective was


FIG. 11

-28 °C, 20.4 bar

-22 °C

New tube


Tube length, m Weight of shell (dry), tons

C2 splitter

From C2 fractionator


LF, 30 fpi 6.5


Polypropylene plant. The first ref-


-6.1 °C FIG. 13

Kettle-type reboiler/condenser with enhanced boiling tubes in heat pump driven C2 splitter. Borealis Polymers ethylene plant in Porvoo, Finland. HYDROCARBON PROCESSING JANUARY 2012

I 59

GAS PROCESSING DEVELOPMENTS TABLE 2. Comparision of LNG chilling train with externally and internally enhanced LF tubes Temperature approach

Externally enhanced LF

Internally enhanced LF




Additional annual income, $MM/yr

0 0 0

16.25 (LNG price = $250/ton)* 26 (LNG price = $440/ton)* 35.8 (.9 LNG price = $550/ton)*

Plot length reduction,%



* Equivalent LNG price in $/MM Btu = 5, 8 and 11 ** Indicates values are compared to the LF (3k) case

equipped with enhanced boiling tubes, with tube OD of 5⁄8 in., was successfully installed in each of the three trains. Test runs following the startup of Train 2 in 2003, Train 3 in 2004 and Train 1 in 2005 confirmed the thermal and hydraulic tube performance. For Train 3, the performance was verified again in 2007, confirming stable performance. In addition, a very low cold approach temperature of 1.4 K between tube-side condensing MR and shell-side boiling propane is confirmed demonstrating the superior performance of the enhanced boiling tube. Qatargas is very satisfied with the overall performance of these chillers. In a joint venture, there are six trains at Ras Laffan, Qatar, with an annual LNG capacity of 7.8 MMtpy per train. All trains are in operation at full capacity. Ethylene plant. Following the first successful application in the polypropylene plant in 2000, further applications followed with various expansion projects and new grassroots projects. Borealis Polymers in Finland, used enhanced boiling tubes in a C2 splitter reboiler/condenser in a heat pump scheme for an ethylene expansion project in 2002. The stable operation has been reviewed and confirmed in 2007, as shown in Fig. 13. Further applications with the enhanced boiling tube followed within the depropanizer and deethanizer condensers, both for the 10th olefin complex for JAM Petrochemical in Iran, and in Yansab, Saudi Arabia. HP LITERATURE CITED Complete literature cited available online at HydrocarbonProcessing.com.

Brigitte Ploix is the manager of Heat Transfer Group, Process and Technology Divi-

L.A. Turbine Europe [email protected] Rue de la Ferme 71- Hall 4 4430 Ans, Belgium Phone: +32 (0) 4 247 30 11


I JANUARY 2012 HydrocarbonProcessing.com

sion, Technip France, Paris, France. She has over 17 years of experience in the thermal design of all nonfired types of exchangers for oil refining and offshore oil production, as well as for the petrochemical, LNG and gas processing industries. Previously, M. Ploix worked as the lead discipline engineer for major international projects and joint ventures. She is a member of the TECHNIP WIELAND Steering Committee, French Association of Oil Industry Engineers and Technicians. Ms. Ploix has been a member of the HTRI Technical Committee since 2008; served as vice chair since 2011; served on the Communication Committee—France since 2003 and chair from 2005–2006. She is a member of the HTRI Plate-Fin Exchanger Task Force. Ms. Ploix is a graduate engineer from the Institut National des Sciences Appliquées de Lyon (INSA,), Lyon, France.

Thomas Lang is the manager of business development for the Process Industry, Product Division High Performance Tubes of Wieland-Werke AG, Ulm, Germany. He has worked Wieland-Werke AG for 19 years. Mr. Lang is responsible for technical marketing and business development for enhanced heat transfer tubes and heat transfer engineering services for the process industry. His experience includes a wide range of enhanced heat transfer application primarily for shell and tube heat exchangers for the oil and gas industry, refining, petrochemical and chemical as well as power industry. Mr. Lang is a member of the HTRI technical committee since 2008 and a member of the HTRI communication committee Germany since 2002. He holds a diploma in mechanical engineering from the University of Stuttgart and an MSc degree from the University of Boulder. 䉳 Select 165 at www.HydrocarbonProcessing.com/RS



Select optimal schemes for gas processing plants Careful process evaluation helps meet product requirements and environmental standards M. MALEKI and M. KHORSAND MOVAGHAR, Energy Industries Engineering and Design Co., Tehran, Iran


as processing plants are an essential part of the energy industry and provide one of the cleanestburning fuels and a valuable chemical feedstock. The importance and complexity of gas processing plants have increased over the years due to their use as a feedstock source and their integration with petrochemical plants. Important factors that drive the process selection and design of gas processing plants are environmental and safety regulations, capital and operating costs, and process efficiency. Therefore, selecting an optimized process scheme during the project feasibility study is vital to ensure that the project is technically feasible, costeffective and profitable.

gas treating of the feed gas using amine is necessary. The amine unit is designed for total H2S removal and total or partial CO2 removal. This article will investigate the reason for partial CO2 removal, which is no longer required due to the recent advances in sulfur recovery technology used in gas processing plants, including South Pars gas plants. Partial CO2 removal scheme. Over

a span of 20 years, the process design of each gas processing plant has been modified to some extent. The main reasons for the changes are more stringent environmental regulations for newer projects and higher ethane purities required for petrochemical plants.

The original scheme designed to meet the product specifications from a sour feed containing high concentrations of H2S, CO2 and mercaptans is depicted in Fig. 1. The gas processing facility has a feed gas treating unit (GTU) using amine for total H 2S removal and partial CO 2 removal. The facility also includes utilities, offsites and necessary infrastructure. The functions of the main process units can be summarized as follows: • Feed reception and gas/liquid separation • Total H2S removal and partial CO2 removal from gas • Dehydration using molecular sieve technology


The process selection study usually begins with a design basis to specify the general configuration of the plant and its outline requirements. These requirements consist of: • Feed characteristics, especially H2S, CO2 and mercaptan concentrations • Product specifications, including maximum concentrations of sulfur and CO2 in the products. The criteria to be optimized for the process scheme selection include: • Environmental and safety compliance with local regulations for effluents from incinerators, flare stacks, wastewater treatment, etc. • Flexibility and performance • Cost • Energy consumption. The maximum H 2 S concentration allowed in the sales gas is 4–5 parts per million by volume (ppmv); therefore,

Flue gas Incinerator Claus unit Sulfur Reception facilities

Gas treating unit (with MDEA)

Gas condensate stabilization

Dehydration unit

Sales gas

Sales gas compressor Ethane recovery unit

Ethane treatment with DEA and drying

NGL fractionation

C3 treatment and drying


C3 liquid

C4 treatment and drying C4 liquid Stabilized condensate FIG. 1

Original scheme for sour gas treating plant.


I 61



• Ethane recovery for production of sales gas and gaseous ethane • NGL separation for production of liquid C3, C4 and gas condensate • Sales gas export compression • Sulfur recovery • Ethane treatment for CO2 removal and drying. Restrictions affect design. New

environmental standards have influenced product specifications and have led to changes in the gas plant design. The main changes that have altered the design are: • Improved sulfur recovery efficiency in the sulfur recovery unit (SRU), resulting in less sulfur being burned in the incinerator • Lower CO2 content requirement in the ethane product

• Lower sulfur content requirement in propane and butane products. This article investigates the best process scheme for the new conditions and compares total CO2 and H2S removal for the GTU vs. partial CO2 removal. It also examines the impact of this change on other units in the plant. Gas treating unit. The GTU uses amine for H 2 S and CO 2 removal. Although we assume that readers are familiar with amines, it is important to note that three types of amines are already used in most gas refineries, as summarized in Table 1.1, 2 When the GTU is designed for total H 2 S and CO 2 removal, normally the CO2 can be reduced to less than 5 ppmv. However, because of design limitations, guaranteed figures using generic dietha-

TABLE 1. Amine technology process capabilities Amine


Regeneration duty

Diethanolamine: (DEA)2HN(CH2CH2OH)2

• DEA is for complete H2S and CO2 removal. • Yields low degradation products and corrosion rates.



• Used selectively for deep H2S removal, with only moderate CO2 removal. • Yields significantly low degradation products and corrosion rates.


Activated MDEA (aMDEA)

• Used for complete CO2 removal; H2S removal also occurs due to enhanced reaction kinetics. • High-performance solvent with reduced energy requirements, high acid loadings, minimal corrosivity and negligible degradation products.


Flue gas Incinerator Acid gas enrichment

Claus unit

Tail gas treatment

Sales gas

Sulfur Sales gas compressor Reception facilities

Gas treating unit (with MDEA)

Dehydration unit

Gas condensate stabilization

Regeneration gas treatment with DEA

Ethane recovery unit

Ethane treatment with DEA and drying

NGL fractionation

C3 treatment and drying


C3 liquid

C4 treatment and drying C4 liquid Stabilized condensate FIG. 2


Sour gas treating plant with partial CO2 removal.

I JANUARY 2012 HydrocarbonProcessing.com

nolamine (DEA) generally provide maximum CO2 concentrations of 100 ppmv in the treated gas. Changes in the sulfur recovery unit. The acid gases (H2S and CO2 ) removed in the amine unit are sent to the SRU. The SRU is based on the modified Claus process for recovering elemental sulfur from acid gas. The chemistry of reactions involved in the Claus process may be described in a simplified form with the following two equations: the first is a simple combustion of one-third of the hydrogen sulfide; the second is the reaction of SO2 produced with the remaining two-thirds of H2S, according to the following reactions:3 H2S + 3/2 O2 t SO2 + H2O – 519 kJ/mole (–124 kcal/mole) 2 H2S + SO2 t 3/n Sn + 2 H2O – 143 kJ/mole (–34 kcal/mole) Feed gas composition. The acid gas from the GTU regenerator column consists of H2S and CO2 that is fed to the Claus unit. The Claus process efficiency is largely dependent on the H2S/CO2 ratio, and it is difficult to maintain at a high value when the H2S content in the acid gas drops below 36%–40%. This difficulty is due to the following factors: • CO2 is an inert gas that dilutes the process gas and, consequently, reduces the conversion efficiency by lowering the partial pressure of reactants. Furthermore, in the reaction furnace (where the thermal Claus reaction takes place), the conversion efficiency is limited due to resulting low flame temperature. • The low H2S content may present challenges to sustaining the flame in the reaction furnace, where only one-third of the acid gas should be burned to achieve the stoichiometric H2S-to-SO2 ratio of 2 required by the Claus reaction. • Carbon dioxide may react with a sulfur species to form carbonyl sulfide (COS) and carbon disulfide (CS2) or a dissociate, which may result in reduction of the overall sulfur recovery, unless adequate precautions are taken for the design of the reaction furnace and the catalytic converters. • Due to the above reasons, limitations in the Claus process on the maximum CO2 content require a ratio of 60:40 for CO2:H2S for the acid gas from the GTU. Lower H2S concentrations can be accepted using special measures (such as enriched air and fuel gas co-firing), but at the expense of lower conversion efficiencies. However, to minimize the Claus unit size and maximize efficiency, an acid gas enrichment using

LNG/GAS PROCESSING DEVELOPMENTS selective amine is generally used to produce “enriched” acid gas suitable for conventional Claus SRUs.4 • Another option is to use an amine absorption process, which is suitable for H2S concentrations within the range of 1%–30 vol%. 5 They offer less complex designs and the same sulfur quality as the Claus process, but have the disadvantage of providing lower sulfur recovery of 96%–97%. Sulfur recovery efficiency. Typical sulfur recovery efficiencies for Claus plants are 90%–96% for a two-stage reactor and 95%–98% for a three-stage reactor. However, new environmental regulations limit sulfur recovery efficiencies to 98.5%– 99.9%. This limitation has led to the development of a large number of tail gas units based on different concepts to remove the last remaining sulfur species.6 Changes in ethane quality. Due to

partial CO2 removal in the current design, the ethane product—which is separated from the treated gas in the ethane-recovery section—contains CO2. Therefore, it becomes necessary to add an ethane decarbonation unit (EDU) to reduce the CO2 level to the specified maximum of 50 parts per million by weight (ppmw) prior to export. DEA is used to remove CO2 from ethane in the EDU before it is dehydrated using molecular sieve beds. In the case of total CO 2 removal in the GTU, the ethane cut from the ethane recovery section will contain very little to virtually no CO2. In fact, ethane will contain around 500 ppmw if CO2 is lowered to 100 ppmv in the GTU using DEA, which can be removed by molecular sieve beds. Also, there is no need for a separate drying unit if molecular sieve beds are used, which provides an advantage in terms of cost, ease of operation and maintenance. Therefore, in the case of total CO2 removal, even the molecular sieve bed is just a CO2 guard. Changes in LPG product quality. The environmental specification for the total sulfur content in LPG products (C3 and C4) has been lowered from 80 ppmw to 10 ppmw; this was the case for LPG products from South Pars gas plants. Sulfur species in LPG are essentially mercaptans, which are removed by direct oxidation with air in the presence of a proprietary catalyst, using a caustic soda-wash process. The LPG product is then dried before export. A molecular sieve sulfur guard is installed after the LPG dryers. The molec-

ular sieve beds are regenerated using sales gas that is then sent to the fuel gas system. The regeneration gas must be CO2-free, as requested by the LPG guard bed vendor. Therefore, if partial removal of CO2 is considered for the GTU design, then another DEA absorber must be installed for total CO2 removal from the regeneration gas. However, if the CO2 is totally removed in the GTU, then molecular sieve guard beds can be used instead of a DEA absorber. Evaluation of existing GTU. The CO2 specifications for sales gas fed to a consumer network should be less than 2 mol%—the maximum limit to prevent general corrosion and pitting in pipelines. This limit will be achieved if the CO2 content in the treated gas from the GTU is less than 1 mol%. In fact, the latter quantity was obtained by back calculation using the ratio of 60:40 for CO2:H2S in the acid gas to the SRU and a sulfur recovery of 97.5%. Therefore, partial CO2 removal in the GTU was dictated by the process requirements in the SRU.


In more recent projects, environmental regulations allow for fewer SO2 emissions from the SRU incinerator, which requires an increase in the conversion of sulfur recovery from 97.5% to 99.5%. Thus, to achieve higher overall sulfur recovery, acid gas enrichment and tail gas treatment units must be used. The acid gas enrichment unit consists of selective H2S removal from the acid gas in the presence of CO2 (i.e., partial CO2 removal from the acid gas). This is accomplished by including an amine unit using a generic MDEA solvent that selectively absorbs H2S. The offgas leaving the top of the amine absorber is sent to the incinerator to convert the residual H2S to less harmful SO2, while the sour gas reaching the required H2S:CO2 ratio from the amine regeneration unit is sent as feed to the Claus reactor. Therefore, by making a gas enrichment unit part of the sulfur gas recovery unit, it is no longer a requirement to specify an outlet CO2 of 1 mol% in the treated gas from the GTU. Based on these results, two alternatives exist for the gas processing plant scheme:

TABLE 2. Effect of CO2 removal on heating value of two sales gases Heating value (Kcal/Nm3) with CO2 (max. 2 mol%)

Heating value (Kcal/Nm3) without CO2

% change

Gas A




Gas B




Flue gas Incinerator

Acid gas enrichment

Claus unit Sulfur

Tail gas treatment Sales gas

Sales gas compressor Regeneration gas

Reception facilities

Gas treating unit (with DEA or aMDEA)

Gas condensate stabilization

Dehydration unit

Ethane recovery unit

Ethane treatment with mole sieve beds

NGL fractionation

C3 treatment and drying


C3 liquid

C4 treatment and drying C4 liquid Stabilized condensate FIG. 3

Sour gas treating plant with total CO2 removal.


I 63



1. Partial CO2 removal in the GTU— similar to the old scheme, regardless of the changes in the SRU, as shown in Fig. 2. 2. Total CO2 removal in the GTU—an optimized gas processing plant scheme based on changes in the SRU, as depicted in Fig. 3. Effects of total CO2 removal on sales gas quality. The sales gas from gas processing plants can contain a maximum of 2 mol% of CO2, as stated in the previous section. Carbon dioxide is an inert gas that only uses energy to be heated to the flame temperature, without any heat input contribution to the combustion. Therefore, its presence at a relatively high amount in the sales gas is a waste of energy. In addition to the above, environmental agencies of many countries continue to implement more stringent emissions standards requiring companies to report their greenhouse gas emissions. Thus, many customers buying sales gas want CO2 levels in the gas to be minimized. Also, the CO2 present in the sales gas is distributed through many users, although it can be recovered for industrial use when totally captured at the gas plant source. These negative aspects of CO2 presence in the sales gas are some of the disadvan-

tages that can be easily prevented by total CO2 removal in the GTU, which can be achieved using DEA or activated MDEA (aMDEA) in the GTU. When the GTU is designed for both H2S and CO2 removal (down to 3 ppmv and 100 ppmv, respectively) in the treated gas using DEA, the CO2 content in the sales gas ranges from 100 ppmv–200 ppmv. Effects of CO2 removal on sales gas heating value. When modifying the sales gas composition by total CO2 removal, it is important to check the heating value to ensure that the changes are not significant enough to require burner change in the consuming furnaces. Two sales gases were studied for this purpose; the results are presented in Table 2. As the table shows, the effect of total CO2 removal on sales gas heating value and its possible consequences on burner design—as well as adverse effects on the operation of existing burners—is insignificant. Effects of total CO2 removal on the dehydration unit. If CO2 is present in the feed gas to the dehydration unit after the GTU, it might be partially co-adsorbed by the molecular sieve beds, resulting in a reduced active area for water adsorp-

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For more information on custom book publishing, event planning or speaking services offered by BIC Media Solutions, contact Earl Heard or Brandy McIntire at (800) 460-4242, email [email protected] or visit www.bicalliance.com/bic-media. For more information on strategic marketing through BIC Alliance, investment banking services through IVS Investment Banking or executive recruiting through BIC Recruiting, contact Earl Heard or Thomas Brinsko at (800) 460-4242, or visit www.bicalliance.com.

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A collection of inspirational & industry stories from business greats of today & yesterday

We Do Together That

Counts & Energy Entrepreneurs

tion and a longer time for bed regeneration. Therefore, it could be expected that, in the case of total CO2 removal, the bed adsorption capacity will be increased while the bed regeneration time and energy consumption are decreased. This is an item that needs further investigation by operators and vendors. Overall optimized scheme. Fig. 3

shows the optimized scheme for total CO2 removal, where the ethane treatment with molecular sieve beds should only be considered if the treated gas from the GTU will provide an ethane product with a CO2 content higher than 50 ppmw. SOLUTION

Before selecting a scheme for a gas processing plant, it is necessary to construct a clear and complete picture of the entire facility. The requirements of each unit within the plant must be understood before they are integrated into the whole scheme. The process scheme selection is carried out during the conceptual stage of a project and should take into account new technology developments for each unit in the plant. Such an approach will deliver an optimized process for the plant that is cost effective, energy efficient, and meets local environmental and safety regulations. HP LITERATURE CITED Complete literature cited available online at HydrocarbonProcessing.com.

Mohammad Maleki is the process, utility and HSE (health, safety and environment) department manager at Energy Industries Engineering and Design Co. He was the consortium process and HSE manager of South Pars front-end engineering and design phases 17 and 18. Mr. Maleki received a BS degree in chemical engineering from the University of Texas at Austin. He has over 30 years of experience as a process and HSE manager, project engineering manager and principal process engineer on several oil, gas and petrochemical projects.

Mohammad Reza Khorsand Movaghar has worked in the process department at Energy Industries Engineering and Design Co. since 2008. He holds a BS degree in petrochemical engineering and a PhD in chemical engineering from Tehran Polytechnic University. Dr. Khorsand also received an MS degree from the University of Science and Technology in Tehran. He served as a process engineer on detailed design projects for gas train units— including acid gas removal, dehydration and ethane recovery—at South Pars gas plant phases 20 and 21. He has over five years of experience as a process engineer and process simulator on several oil and gas projects.



Improve process control for natural gas heat exchangers Dynamic simulation model identifies how to optimize plant controllability and safety H-M. LAI, Jacobs Canada Inc., Calgary, Alberta, Canada


ynamic simulation is becoming an important tool for engineering design and plant operation.1,2 In this case history, a first-principle dynamic model of a natural gas/steam heat exchanger system is built using a commercially available dynamic simulator. Four scenarios for operability and safety are investigated to demonstrate how a process and associated control system will respond to various disturbances as a function of time.

Case history. Preheating of natural gas (NG) is frequently used to prevent hydrate formation due to the Joule-Thompson effect of the NG let-down stations. The typical NG heater system consists of a shell-tube heat exchanger, a condensate receiver and a steam trap, as shown in Fig. 1. The steam control arrangement is also shown in Fig. 1. This system includes: • One temperature control valve on the steam inlet line • One level control valve on the condensate outlet line • One pressure control valve on the vapor outlet line. As a part of the plant design, a steadystate simulation of the system is done to check the heat-and-material balances and equipment sizing. Table 1 lists the process conditions and major equipment sizing data.

Fig. 2 shows the process response to an inlet temperature change of the NG from 0°C to 10°C. As shown in Fig. 2, when the NG inlet temperature rises from 0°C

LP steam

to 10°C as a result of falling heat load, the steam pressure in the heat exchanger will drop about 200 kPa. Fig. 3 shows the process response to a change in NG











VOutA Val-A

ToDrumA 260MC22002A



LiqA 260MG22002A

LiqOutA POutA PLiqA

PumpE FIG. 1

Simplified process scheme for the NG steam heater.

FIG. 2

Process responses to NG inlet temperature changes: 0°C to 10°C.

Scenario 1: Process upsets. In this

scenario, the impacts of both inlet NG temperature changes from 0°C to 10°C and NG demand changes from 100,000 kg/h to 140,000 kg/h have been analyzed. The simulator logic unit operation—the transfer function block—is used to simulate sine wave changes of NG demand and inlet temperature.


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LNG/GAS PROCESSING DEVELOPMENTS flow from 100,000 kg/h to 140,000 kg/h. Note: The steam pressure also drops over 250 kPa, while the NG demand declines from 140,000 kg/h to 100,000 kg/h. These dynamic simulation results confirm that steam pressure in the condensate receiver cannot be maintained at stable ranges during process upsets. If a steam trap is used, then the steam control scheme will lead to reduced condensate flow from the steam heater system, and it will form the so-called “stall behavior.” Scenario 2: Stall behavior. This sce-

nario discusses condensate removal from the heat exchanger. As mentioned before, the temperature control valve on the steam line maintains the NG outlet temperature by opening or closing to adjust the steam flowrate, thereby varying the steam space pressure. When the steam pressure in the heat exchanger is equal to, or less than, the total backpressure imposed on the steam trap, then the reduction or cessation of condensate flow from the heat exchanger occurs. The condensate will back up in the drain line and will flood back into the exchanger. This condition can damage the control valve and may cause corrosion of the exchanger. This symptom is called the “stall behavior.” Based on current heat exchanger sizing data, the dynamic heat model for this steam heater system was built using dynamics and spreadsheet tools. These conditions were assumed for the model: • NG gas inlet temperature rising to 10°C

FIG. 3

Process responses to NG demand changes.

FIG. 4

Simulated results of stall behavior for the steam heater.

TABLE 1. System design data and operating conditions LP steam JTIC BPOut

Bypass NG





JValve NGOutA






VOutA Val-A

ToDrumA 260MC22002A


LiqA 260MG22002A PumpE FIG. 5


Alternative control scheme for the NG steam heater.


LiqOutA PLiqA

Steam inlet pressure, kPaa


Steam inlet temperature, °C


Steam flowrate, kg/h


Natural gas inlet pressure, kPaa


Natural gas inlet temperature, °C


Natural gas flowrate, kg/h Natural gas density @ OP, kg/m3


Ratio of specific heat


Steam exchanger type




Total surface area, Overall

U, W/m2


Tube inside diameter, In Condensate receiver size


I JANUARY 2012 HydrocarbonProcessing.com


800 0.834 1m ID x 3m T/T

LNG/GAS PROCESSING DEVELOPMENTS • Low-pressure (LP) steam pressure of 442 kPa and steam-trap backpressure of 338 kPa • NG consumption of 132,000 kg/h. Fig. 4 illustrates the simulated stall behavior. Due to the 10% over design margin, the heat exchanger has more heating area than required. So, the operating steam pressure will be much lower than needed. When the condensate is waterlogged in the heat exchanger, the surface area available to condense steam is reduced. The heat flow drops, and NG outgoing temperature begins to fall. While the temperature sensor detects this change, the controller will open the steam control valve. This raises the pressure in the steam space to above the trap-back pressure and causes condensate to pass through the trap. The condensate level falls, and the NG temperature climbs. When the sensor detects this, the controller closes the control valve. The steam pressure falls, and then flooding begins again. The result is a continual cycling of opening and closing the steam control valve. The side effects of stall include damaging the control valve and water hammer along with corroding and leaking heat exchangers. These operating conditions will increase maintenance incidents and reduce the service life of the steam heater and associated equipment.

7. The results show that the maximum change of steam pressure is much less, lower than 30 kPa. The stall behavior will


not happen, as the pressure in the steam space is always greater than steam trap backpressure. Compared with the regular

FIG. 6

Process responses to NG inlet temperature changes: 0°C to 15°C.

FIG. 7

Process responses to NG demand changes.

FIG. 8

Tube rupture profiles for 4M6 PSV with 6 in. inlet/10 in. outlet piping.

Scenario 3: Alternative control scheme. There are different ways to

prevent stall.3 Normally, we could use an alternative means to remove condensate from the exchangers by installing a pumping trap, instead of using steam traps if the pressure in steam space may be less than the backpressure. We could also size the heat exchangers and steam traps properly to ensure that the pressure in steam space is stable and always higher than the backpressure under all operating conditions. Or we should reduce the backpressure of condensate discharge lines. In reality, this can’t always be done. For the present NG steam heater, the most cost-effective solution is to use an alternative control scheme—a bypass control. This control approach bypasses a partial NG stream around the exchanger and blends it with a fraction that has passed through, as shown in Fig. 5. The temperature control valve is relocated from the original steam line to the NG bypass line. System dynamic responses to the process upsets over NG demand and inlet temperature are illustrated in Figs. 6 and


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FIG. 9


Tube rupture profiles for 4P6 PSV with 8 in. inlet/12 in. outlet piping.

TABLE 2. PSV sizing results for tube rupture

the tube-rupture event. After about 3 seconds, the receiver pressure reaches the set pressure; then the PSV starts to relieve. • PSV would work fine if the normal PSV of 4M6 sized by a conventional method is installed on the top of shell side in the steam heater. However if this PSV is relocated to the top of the condensate receiver, a 40% overpressure in the receiver would occur, as shown in Fig. 8. The major reason is that, under the upset conditions of the tube rupture, the NG has a strong stripping effect (due to vapor/liquid equilibrium) that carries the steam out of the condensate phase. This causes the PSV peak relief load (30,530 kg/h) from the condensate receiver to be about 23% higher than the tube-rupture flow (24,920 kg/h) estimated by API 521 method. • As evaluated in Fig. 9, if installing a PSV on the top of the condensate receiver, a larger sized 4P6 PSV and associated larger inlet/outlet piping should be installed. This example shows that, when upset conditions occur, equilibrium conditions in vessels are changing, and the safety system design must be adjusted to account for those changes.



Shell side

Top of receiver



PSV type



PSV sizes



PSV set pressure, kPaa



PSV full open pressure, kPaa



Options. This case study illustrates how

PSV closing pressure, kPaa



critical it is to consider vapor/liquid equilibrium changes and interaction of process with controls in the system design, and how dynamic simulation can improve plant performance, controllability and safety in design and operation. HP

PSV installed location Relief load, kg/h

Inlet piping size, in.



Inlet loss % of set pressure



Outlet piping size, in.



Outlet loss % of set pressure



steam control, the simple bypass control greatly improves the operating performance of the steam heater. Scenario 4: Tube rupture contingency. Pressure-relief systems are a critical

part of any process design. Proper design of these systems is required by regulation and industrial codes. Due to the large operating pressure difference between the exchanger tube and shell sides (flange rating 900 lb at tube side vs. 150 lb at shell steam side), the case of complete tube rupture is a valid case in the steam heat exchanger. Although the simulator cannot predict the instantaneous pressure wave at the rupture site, it does provide important insights on the dynamic system behavior under the tube-rupture conditions. Normal operating data and pressure safety valve (PSV) sizing results by the conventional method are listed in Tables 1 and 2. These parameters were set to generate the 68

I JANUARY 2012 HydrocarbonProcessing.com

initial values of the dynamic model for the tube-rupture case: • UA value was set for the steam heat exchanger • Condensate receiver was set to real sizes to simulate steam/liquid accumulation and liquid level variations • Normal valve with a customized spreadsheet was used for constant NG rupture flow into the steam condensate system. In general practice, to protect overpressure of steam system from the high pressure of NG, a check valve should be installed on the upstream steam line, and a PSV shall be provided on the top of the vapor line in the condensate receiver. The dynamic simulation with two different PSV sizes was verified, Figs. 8 and 9 summarize the results. Some highlights of the dynamic simulated results are discussed here: • Pressure in the condensate receiver begins to build up immediately following

ACKNOWLEDGMENTS Special thanks to Alan Childs, manager of the process department, for the valuable discussions, review and comments. LITERATURE CITED Dissinger, G. R., “Studying simulation,” Hydrocarbon Engineering, May 2008. 2 James, G. and J. Reeves, “Dynamic Simulation Across Project and Facility Lifecycles,” 6th World Congress of Chemical Engineering, Melbourne, Australia, Sept. 23–27, 2001. 3 www.spiraxsarco.com/Resources, “Practical Methods of Preventing Stall.” 1

Hai-Ming Lai is a principal process engineer in Jacobs Canada Inc., Calgary, Alberta, with over 26 years of experience in process research and development, design, and engineering of oil and gas, refining/upgrading, and petrochemical projects. His specialties include simulation studies, conceptual and front-end engineering design. He holds a PhD in chemical engineering from Beijing University of Chemical Technology (BUCT), P.R. of China., and is a registered professional engineer in Alberta, Canada. Prior to joining Jacobs, Dr. Lai worked for Aspen Technology, Calgary, Canada, and Research Institute of Chemical Technology in BUCT, Beijing, P. R. of China.


Consider lobe blowers combined with compressors New blower meets low-pressure applications cost-effectively H. P. BLOCH, P.E., HP Staff


ir and inert gases are often compressed in a variety of processes and facilities. These processes and applications include wastewater treatment plants, oil refineries, the petrochemical processing industry and power generation plants. Compression equipment is also installed where a carrier gas such as air is needed for conveying powders. In general, these applications require gas pressures that are too low for what we typically call a compressor and too high for what is usually handled by blowers. Of course, these services could always be satisfied with independent or free-standing machines, but promising hybrid machines are now available.

New developments. Recently, a particular hybrid machine that combines both blower and compressor technologies has joined the list of available options (Fig. 1). It is called a lobe blower compressor. This blower compressor draws on both technologies and bridges an important gap. Following extensive field tests under harshest operating conditions and after over three years of successful operation in various fields of endeavor, these hybrid packages are now finding much wider markets.

that combines the essential concepts of lobe blowers and compressors. One well-known hybrid manufacturer relies on more than 100 years of experience designing and producing rotary lobe blowers. The same company has produced thousands of process screw compressors in the decades since 1943. The performance envelope of its single-input hybrid gas mover opens new possibilities for producing either positive pressure or vacuum. High-energy efficiency with lower LCCs. The goals for process

machinery are summarized in lowest lifecycle cost (LCC). The concept aims for and takes into account highest energy efficiency, infrequent maintenance, and low risk of unscheduled downtime. A complete package, low noise levels without the need for additional soundproofing, space saving, operating ease and a wide operating range are desirable attributes of these machines. Over a 10-year operating period, energy costs equate to 90% of the total LCCs of a compressor. The actual equipment costs play a secondary role (Fig. 3). With this in mind, one such hybrid gas mover was

Lobe blower compressors cover wide performance range. Low-

pressure (LP) applications often use the roots-principle of isochoric (constant volume) compression in a rotary lobe arrangement. In contrast, screw compressors that follow internal volume reduction principles become the preferred choice because of their relatively high-energy efficiency in the elevated pressure ranges. The latest hybrid design package (Fig. 2) goes beyond the standard rotary lobe blower and screw compressor designs. Both blower and screw compressor have been optimized and upgraded in this hybrid

recently developed with the focus on increased energy efficiency and reduced greenhouse gas emissions. Compared with some stand-alone rotary lobe blower and screw compressor technologies, this innovative hybrid was reported to reduce energy consumption by up to 15%. The screw (also known as rotary-lobe) compressor stage benefits from rotor pro-

FIG. 2

Package design incorporating a rotary lobe—screw compressor hybrid.

90% Energy

5% Maintenance 5% Initial cost

FIG. 3 FIG. 1

A new hybrid design incorporates rotary lobe blower and screw compressor principles (Source: Aerzen USA, Kulpsville, Pennsylvania, US).

Average operating costs of an air mover over 10 years- Largely composed of energy, initial cost, maintenance: 90% Energy (white), 5% maintenance (blue) and 5% initial cost (red).


I 69

ROTATING EQUIPMENT file innovations and a low-loss inlet cone. Together, these innovations produce important efficiency gains. Optimal air flow within the acoustic enclosure package directs cool air to the intake side and increases compression efficiency. The equipment package incorporates special silencer insulation. High-volume flow models were tested to have control ranges from 25% to 100%. Innovation was also noted in other ways. New sealing solutions are used at the drive shaft and at the rotor chamber to minimize seal wear. Rolling element bearing configurations were selected and designed to extend bearing L-10 life to over 60,000 operating hours at a differential pressure of 1,000 mbar (approximately 15 psi). Oil change intervals of 16,000 operating hours are now entirely feasible. As has been customary for sensitive services in the past, a hybrid should routinely feature a purely reactive discharge silencer. Since absorption material breaks down over time, knowledgeable hybrid manufacturers will not use these materials so as not to contaminate a downstream process system. This is important in pneumatic conveying systems for bulk

materials in the food industry. Reactive discharge silencers will avoid accumulation of broken-down absorption material in the fine diffuser systems of wastewater treatment plants, which is helpful from a maintenance avoidance point of view. In essence, preventing clogging saves the cost of cleaning and possibly reducing plant operating capacity. How new hybrids expand flow and pressure range capabilities.

By design, conventional rotary lobe blowers are limited to a differential pressure of 1 bar (~15 psi). When it comes to higher pressures, other types of compressors are used. These are often designed for significantly higher pressures and at higher initial investment. As was already mentioned, hybrids bridge the gap. One very recently developed hybrid is called a rotary lobe compressor. It incorporates a differential pressure capability increased to now 1.5 bar (22 psi). Vacuum operation is feasible in the extended range from 500 mbar (–1 5 in. Hg) to –700 mbar (–21 in. Hg). The discharge temperature limit of this modern rotary lobe blower-compressor combination has been increased

from previously 160°C (320°F) to now 180°C (356°F). One typical line of hybrid rotary lobe compressors covers a flow range from 110 m³/h to 4,100 m³/h (65 cfm to 2,400 cfm) with 12 machine sizes. Its controls are based on modular design concepts that can be tailored to each individual application. These controls incorporate a frequency converter and power supply panel ready to be plugged into existing power supplies and plant piping. In modern hybrid controls, all measured operating data are retrievable and parameters adjustable from user-friendly keypads. HP Heinz P. Bloch is a consulting engineer residing in Westminster, Colorado ([email protected]). He has held machinery-oriented staff and line positions with Exxon affiliates in the US, Italy, Spain, England, The Netherlands and Japan in a career spanning several decades prior to retirement as Exxon Chemical’s regional machinery specialist for the USA. Mr. Bloch is the author of 18 comprehensive texts and close to 500 other publications on machinery reliability improvement. He advises process plants worldwide on equipment uptime extension and maintenance cost-reduction opportunities. He is an ASME Life Fellow and maintains registration as a professional engineer in Texas and New Jersey.

The industry-standard software for instrumentation design Featuring more than 70 routines associated with control valves, rupture disks, flow elements, relief valves and process data calculations, InstruCalcTM is one of the industry’s most popular desktop applications for instrumentation calculations and analyses. Features: NEW • Graphs for Control Valves and Flow Elements Version 8.1 • Restriction devices • Material yield strengths file • ISO orifice plate calculations have been updated to ISO 5167, 2003 sudden entrance and exit to the calculations. • Relieff VValve alve ve pprograms, ve rg ro +1 ((713) 520-4426 l [email protected] +1 om www.GulfPub.com

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SERVING YOU WITH PRIDE THE 2012 AFPM (NPRA) ANNUAL MEETING Don’t miss your industry’s most important gathering of the year! Manchester Grand Hyatt Hotel San Diego, California March 11 – 13, 2012 afpm.org

In 2012, we’re changing our name from NPRA to AFPM, American Fuel & Petrochemical Manufacturers to better describe who we are and what we do. And you can still count on us to bring you the most comprehensive conference program for the refining and petrochemical industries. Select 97 at www.HydrocarbonProcessing.com/RS

This year’s program has something for everyone from process and operations engineers to senior management. Keynote sessions feature top government officials and corporate CEOs. Breakouts cover nearly every facet of refining technology. And networking opportunities abound at our receptions and affiliate-hosted events. Help us celebrate our new beginning. Register at www.afpm.org or www.npra.org.


Are you losing money when tuning controllers? Here are 10 rules, if followed, that will result in poor process performance M. J. KING, Whitehouse Consulting, Isle of Wight, UK


rom published literature, there is a wide range of pitfalls into which control engineers frequently stumble.1–3 As these pitfalls were documented, more were discovered. In this article, we will focus on a very specific area of control design. The following rules investigate how substantial deterioration in process performance is possible in proportional, integral and differential (PID) control systems.

Rule 1. Use the ‘derivative-on-error’ algorithm. The

PID algorithm in its conventional analog form is usually written as:

E ×dt +Td

Rule 2. Use the ‘proportional-on-error’ algorithm.

Using this algorithm is almost entirely to blame for hiding opportunities to substantially improve the performance of controllers responding to process disturbances. The alternative “proportionalon-PV” offered as an option in most DCS is described as: ⎡ ⎤ ts ⎢(PVn − PVn−1 ) + E n + ⎥ ⎢ ⎥ Ti ⎥ ΔM = K c ⎢ ⎢Td ⎥ ⎢ ( PVn − 2PVn−1 + PVn−2 )⎥ ⎢⎣ ts ⎥⎦

dE ⎤⎥ dt ⎥⎦

Despite this or, more often, its equivalent in Laplace form, being used in most distributive control systems (DCSs) vendors’ documentation it strictly applies only to analog control. A close digital equivalent is: ⎡ ⎤ T ts ΔM = K c ⎢(E n − E n−1 ) + E n + d ( E n − 2E n−1 + E n−2 )⎥ ⎢ ⎥ ts Ti ⎣ ⎦

The problem with this algorithm is that when the setpoint (SP) is changed, assuming the process was previously at steady state, the derivative action causes an immediate step change in output, given as: K cTd ΔSP ts This is followed, at the next scan interval, by the same change in the opposite direction. Known as the “derivative spike,” it can readily move the manipulated variable (MV) full scale. Td might typically have a value of around 1 minute, and ts will be about 1 second. Even with quite a modest value for Kc , ΔM can exceed 100%. Fortunately, most DCS vendors have modified the algorithm to: ΔM =

⎡ ⎤ T ts ΔM = K c ⎢(E n − E n−1 ) + E n + d ( PVn − 2PVn−1 + PVn−2 )⎥ ⎢ ⎥ Ti ts ⎣ ⎦

Known as the “derivative-on-PV” algorithm, the derivative action no longer responds to changes in SP. However, the response to changes in process variable (PV), caused by process disturbances (or “load” changes), is unaffected. Some DCS vendors have retained the derivative-on-error version as an option—unfortunately, often as the default version. A poorly trained engineer

At first glance, this might appear to have a serious disadvantage. When the SP is changed, the more conventional proportional-on-error algorithm generates a “proportional kick” equal to Kc ΔSP—doing much to ensure that the SP is approached rapidly. The proportional-on-PV version does not do this, relying entirely on the much slower integral action. Many engineers reject this algorithm solely because of this perceived problem. However, they overlook the fact that the controller can be re-tuned to compensate for the loss of the proportional kick. As shown in Fig. 1, with effective tuning, its response to SP changes would be virtually indistinguishable, by the process operator, from that of the algorithm it replaces. Its benefit becomes clear when the performance of the two algorithms, both tuned for SP changes, is compared for load


⎡ 1 M = K c ⎢E + ⎢ Ti ⎣

might think that, since it bears the closest resemblance to the conventional analog version, it should be the one to apply. This seriously limits the use of derivative action in those situations where it would be particularly beneficial (See Rule 7).

SP Proportional-on-error Proportional-on-PV Time

FIG. 1

Response to a SP change.


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PROCESS CONTROL DEVELOPMENTS changes. With the same tuning, provided the SP remains constant, the two algorithms perform identically. The much faster tuning necessary to make the proportional-on-PV algorithm perform well for SP changes causes it to respond much faster to load changes. Fig. 2 shows that both the duration of the disturbance and the maximum deviation from SP are typically halved. Were the PV to be related to product composition, the volume of off-spec production would be reduced by more than 75%. Of course, it would be possible to achieve the same improvement by applying the tuning developed for the proportional-onPV algorithm to the proportional-on-error version. However, it would then cause a major process upset whenever the SP is changed. This perhaps explains why the algorithm is not fully appreciated. Many engineers select the more conventional proportional-on-error algorithm and tune it for SP changes. Its response to load changes will then appear reasonable but will disguise the fact that the response can be substantially improved. Rule 3. Use the interactive algorithm. There is an alter-

native derivation of the PID controller. It starts with a conventional PI controller, but adds the derivative action by replacing the E term with a “projected error” defined as: dE Eˆ = E +Td dt

This results in a slightly different algorithm: ⎡⎛ T ⎞ 1 dE ⎤⎥ M = K c ⎢⎢⎜⎜1+ d ⎟⎟⎟ E + ∫ E . dt +Td Ti dt ⎥⎥⎦ Ti ⎟⎠ ⎢⎣⎜⎝ Comparison with the so-called “ideal” form described earlier shows that the integral and derivative actions are unchanged but the proportional action depends not only on Kc but also now on Ti and Td —thus earning the algorithm its “interactive” name. Some DCS use this version, either as the only choice or as an option. It exists primarily because it closely matches the action of pneumatic analog controllers and their early electronic replacements. Using it these days presents no problem provided the tuning method chosen is specifically designed for the changed algorithm. Indeed, provided that in the ideal algorithm Td is less than 0.25 Ti, it is possible to calculate equivalent tuning for the interactive version so that the performance of the two algorithms is identical. And if the derivative is not used, then both algorithms are the same in any case. The problem arises because DCS vendors rarely retain the algorithm in its pure form. It is common to include a “derivative filter” (usually given the nomenclature as a or ␣) or a “derivative


SP Proportional-on-error Proportional-on-PV


FIG. 2


Response to a load change.

I JANUARY 2012 HydrocarbonProcessing.com

gain limit” (which is the reciprocal of a). This value may be fixed within the system or configurable by the engineer. It usually makes impossible adapting a tuning method designed for the ideal algorithm for use with the interactive form. Rule 4. Apply Ziegler-Nichols tuning. Amazingly, Ziegler-Nichols is still by far the most popularly taught tuning method. It was developed 70 years ago.4 Few appreciate that it assumes the now rare interactive version of the PID algorithm. Even fewer know that it was developed for load changes and so, if applied to the normal proportional-on-error algorithm, will result in far too an aggressive response to a change in SP. And, even if these issues are resolved, its main objective is to deliver the “quarter decay ratio,” where the height of the second PV overshoot is one quarter of the height of the first. Few now accept that any amount of second overshoot is the sign of a well-tuned controller. The more cynical control engineer might think inclusion of the method in papers and textbooks is to establish a benchmark by which even a poorly performing alternative would look good. Another commonly reproduced method is that developed by Cohen-Coon.5 It too uses the quarter decay ratio and was developed using analog control almost certainly equivalent to the interactive algorithm. If anything, its performance is somewhat inferior to Ziegler-Nichols. Rule 5. Ignore the MV. Effective controller tuning is often a compromise between a fast return to SP and avoiding excessive changes to the MV. Many tuning methods use a penalty function, such as the integral over time of absolute error (ITAE), as a measure of control performance: ∞

ITAE = ∫ E t. dt 0

Minimizing such functions results in the fastest possible return to SP but, if the deadtime-to-lag ratio is small, this will result in excessive adjustments to the MV. As the deadtime-to-lag ratio approaches zero, such methods recommend a controller gain approaching infinity. One such method is that developed by Smith, Murrill and others.6,7 Defining the MV overshoot as the percentage by which the peak change in MV exceeds the necessary steadystate change, we can supplement this type of tuning criterion by minimizing the penalty function subject to a limit on MV overshoot. Typically, a 15% limit results in what most would accept as a well-tuned controller. However, the limit may be increased if large changes in MV do no harm and similarly reduced if the aim is to minimize MV movement. Indeed the latter, in the case of surge vessel level control, is the overriding consideration, and large deviations from level SP should be the norm. One of the few published methods that permits the engineer to specify the compromise between fast return to SP and MV movement is internal model control (IMC) tuning. Several companies have adopted this method as standard. However, it does have a number of disadvantages. The method is derived using “direct synthesis,” which develops a control algorithm that will respond to an SP change according to a defined trajectory. This is usually specified as an approach to SP with a user-specified lag of ␭. The resulting tuning equations vary greatly. For example, it can be applied to both self-regulating and integrating processes, using either the ideal or interactive algorithm. The synthesis usually includes terms that are not part of the PID algorithm and, so, some approximation is necessary or the terms simply

PROCESS CONTROL DEVELOPMENTS ignored. Different developers reach different conclusions. But a common example for the ideal PID algorithm applied to a selfregulating process is: θ τ+ θ τθ 1 2 Kc = Ti = τ + Td = 2 2τ + θ K p λ +θ While the method permits the user to decide how aggressive the control should be, the value of ␭ has to be determined by the trialand-error method. While some texts provide some guidance, there is no predictable relationship between its value and MV overshoot. Under a different set of process dynamics, the relationship changes. It is possible to develop formulae for the best choice of ␭. For example, choosing a value given by 0.31␪ + 0.88␶ will give an MV overshoot of 15%, but only for the proportional-on-error form of the ideal controller applied to a self-regulating process. We would need to develop such formulae not only for different controllers and for integrating processes but also for different MV overshoot limits. While perhaps possible, the most damning limitation of this tuning method is that no one has yet published the formulae for the preferred algorithm—where both proportional and derivative actions are based on PV rather than error.

with either little or a large deadtime—depending on the disturbance source. Fig. 4 shows the impact on ITAE of removing deadtime from a well-tuned controller, and retuning the PI controller as well as possible. It shows that for SP changes, removing derivative action causes controller performance to deteriorate more on processes that have a larger deadtime-to-lag ratio. For load changes, the opposite is true. But, for both cases, the effect of removing it is always adverse, and, in any case, most controllers have to deal with both disturbance types. In practice, the derivative action is only used by a minority of controllers. There are several reasons for this. First, it has a reputation for causing problems if there is measurement noise. Certainly, it will grossly amplify noise, but modern DCSs do offer a wide range of filtering techniques that can readily reduce noise to a point where derivative action is viable. Second, it adds another tuning parameter. Adding derivative action requires the proportional and integral tuning to be readjusted. Fig. 5 shows that the addition of derivative action is beneficial because it permits a larger controller gain. If the engineer has already spent hours tuning a PI controller by the trial-and-error method, there will be an understandable reluctance to abandon this tuning and start afresh with a three-dimensional search.

Rule 6. Ignore the scan interval. The industry has now

Rule 8. Use filters to improve PV trending. Most con-

begun replacing first generation DCSs with their more modern counterparts. Engineers have been surprised to find in some cases that this has apparently increased the level of measurement noise. This can arise because of the faster scanning that may be available in the new system. Fig. 3 shows how the total valve travel generated by a PID controller varies as scan interval changes. The curve starts at a ts/␶ ratio of 1/120—equivalent to a controller with a scan interval of 1 second on a process with a lag of 2 minutes. Defining the total valve travel under these conditions as 100%, we can see that, for a PID controller, reducing the scan interval from 2 seconds to 1 would increase valve travel by a factor of 4. All DCS include the ability to filter a measurement and most use the first order exponential type. The digital version of this filter is often defined as:

trol engineers use filters to make the PV trend look good. Gone are the days when we have to concern ourselves with the amount of ink used in drawing such trends. A better criterion is to examine the movement of the final actuator, usually a control valve. This will depend not only on the amplitude of the measurement 100

∑(ΔM ), %






Yn = P ×Yn−1 + (1− P )X n

Changing the scan interval of a controller in a system in which the engineer defines P directly will result in a different filter lag. Even the most modern of controller tuning methods still assumes analog control. While this is of little concern when the scan interval is small compared to the process dynamics, it can cause problems otherwise. For example, compressor-surge protection systems are applied to a process where the deadtime is effectively close to zero and the lag only a few seconds. Tuning such controllers without taking account of scan interval will drastically affect performance. It goes some way to explain why package vendors (usually mistakenly) insist that compressor controls can only be implemented in special purpose control systems that have a much shorter scan interval.

0.01 0.0

FIG. 3







Effect of scan interval on noise passed to an actuator.

300 250 Increase in ITAE, %

where P = exp(−ts / τ f )

200 150 100

Load change


SP change


Rule 7. Avoid using derivative action. Depending

on the textbook a control engineer might read, if the process has a large deadtime, the derivative action is either beneficial or becomes less effective. In fact, it offers an advantage on processes

0 FIG. 4






Impact in ITAE of removing derivative action.


I 75

PROCESS CONTROL DEVELOPMENTS noise but also on the controller tuning. If the impact on valve movement is acceptable, then the filter serves no purpose and will reduce the controllability of the process. Its presence means that tuning has to be relaxed to maintain stability. Conversely, we must remember that, if a filter is removed, then the benefit will not be apparent until the controller is re-tuned to accommodate the change in apparent process dynamics. Filtering can be beneficial if it permits greater use of derivative action. Since derivative action responds to the rate of change error, the small fluctuations in signal occurring at a high frequency are greatly amplified. Many DCSs now offer the facility to selectively filter only the measurement passed to derivative action. This permits derivative to be used without changing the dynamics seen by the proportional and integral actions. Rule 9. Tune by trial-and-error methods. Over 200 tuning methods have been published.8 All of them have at least one flaw. It is not surprising that control engineers have generally adopted the trial-and-error method as the main tuning method. It requires no knowledge of the process dynamics and little understanding of the control algorithm being applied. But its main disadvantage is that it is extremely time-consuming. Trials conducted on a simulated process with dynamics of a few minutes showed that engineers would spend around 30 minutes finding the best tuning. Quite a modest investment one might think until one realizes that the simulation was running much faster than real time and each test was exactly reproducible. On the equivalent real process such an exercise would easily have filled a working week.

In practice, no engineer can commit this time to a single controller and will stop trying to improve its performance once it is stable and looks “about right.” The result is that the process operator will likely be unimpressed by its performance during the next process upset and will switch the controller to manual. Developers of tuning methods have attempted to develop a set of tuning formulae that can be applied to any situation. In reality, such an approach is unlikely ever to be successful. There are two fundamentally different processes: self-regulating and integrating. There are two fundamentally different PID algorithms: ideal and interactive. Some versions of the algorithm include a derivative filter that cannot be changed by the user. Proportional action can be based on error or PV, as can derivative action. These options are not mutually exclusive; just considering those listed so far gives 32 possible combinations. If we add to this the requirement to specify the aggressiveness of the control, allow for different scan intervals and to take account of vendor-specific modifications to the algorithm, then the number of sets of tuning formulae grows to an impractical level. Figs. 6–8 show comparisons between the commonly published tuning methods and user-defined optimum tuning. For the comparisons to be fair, the controller was assumed to be analog and subject to a SP change. The results were obtained by using a tuning constant optimizer freely available.9 In this case, the optimum tuning was specified as minimum ITAE subject to a 15% MV overshoot limit. So, unlike many methods, the optimized controller gain does not approach infinity as ␪/␶ approaches zero. The IMC method appears to estimate the con3





Ti /τ

log10 (Kp · Kc)




-0.1 PI

-0.3 -0.5 0 FIG. 5















Impact on Kp of inclusion of derivative action.

FIG. 7

Determination of integral time.




1.0 0.8 Td /τ

Kp ·Kc

Optimized (P on E) Optimized (P on PV) λ = 0.31θ + 0.88τ Smith, Murrill et al. Cohen Coon Ziegler-Nichols

0.6 0.4 0.2

0.1 0



ϑ/τ FIG. 6


Determination of process gain.

I JANUARY 2012 HydrocarbonProcessing.com



0.0 0



ϑ/τ FIG. 8

Determination of derivative time.

PROCESS CONTROL DEVELOPMENTS troller gain well, but only because the choice of ␭ has been optimized for this particular case. Note: The method developed by Smith, Murrill and others is only applicable to values of ␪/␶ less than 1. Outside of this range, it can generate negative tuning constants. But, most importantly, optimization permits tuning to be derived also for the preferred proportional-on-PV algorithm. The much higher gains derived for this controller will substantially reduce the impact of process disturbances. Rule 10. Don’t train engineers in basic control. The

most effective way of reducing process profitability is to ensure that the control engineers are kept completely unaware of what can be achieved by minor changes to PID control. Those that have studied control theory at university will have been subjected to daunting mathematics, much of which is irrelevant to the process industry. Almost certainly little will have been covered on the alternative forms of the PID algorithm, let alone which one to use and how to properly tune it. While it is common practice to send staff on vendor supplied courses in DCS programming and multivariable predictive control (MPC), it is rare to consider also training in basic control techniques. Industry seems to expect engineers to somehow acquire this expertise without outside assistance. This ensures that the techniques described above, many of which have been available for over 30 years, are still not properly appreciated and that plants continue to operate away from maximum profitability. HP NOMENCLATURE Complete nomenclature available online at HydrocarbonProcessing.com.

LITERATURE CITED King, M. J., “How to lose money with advanced controls,” Hydrocarbon Processing, June 1992, pp. 47–50. 2 King, M. J., “How to lose money with basic controls,” Hydrocarbon Processing, October 2003, pp. 51–54. 3 King, M. J., “How to lose money with inferential properties,” Hydrocarbon Processing, October 2004, pp 47–52. 4 Ziegler, J. G. and N. B. Nichol,“ Optimum settings for automatic controllers,” Transactions of the ASME, 64, pp. 759–768, 1942. 5 Cohen, G. H. and G. A. Coon, “Theoretical considerations of retarded control,” Transactions of the ASME, 75, pp. 827–834, 1953. 6 Smith, C. L., Digital Computer Process Control, Intext Educational Publishers, p. 147, 1972. 7 Lopez, A. M., J. A. Miller, C. L. Smith and P. W. Murrill, “Tuning controllers with error-integral criteria,” Instrumentation Technology, 14, pp. 57–62, 1967. 8 O’Dwyer, A., “A summary of PI and PID controller tuning rules for processes with time delay,” IFAC Digital Control: Past, Present and Future of PID Control, Terrassa, Spain, 2000. 9 http://www.whitehouse-consulting.com/tune.htm. 1

BIBLIOGRAPHY King, M., Process Control: A Practical Approach, published by Wiley, ISBN 9780-470-97587-9.

Myke King is the author of Process Control: A Practical Approach. He is the director of Whitehouse Consulting. Previously, he was a founding member of KBC Process Automation, and prior to that he was employed by Exxon. He is responsible for consultancy services assisting clients with improvements to basic controls and with the development and execution of advanced control projects. He has 35 years of experience in such projects, working with many of the world’s leading oil and petrochemical companies. Mr. King holds an MS degree in chemical engineering from Cambridge University and is a Fellow of the Institute of Chemical Engineers.

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How to manage vaporization in an analytical system When done properly, this process ensures that all compounds vaporize at the same time, preserving the sample’s composition D. NORDSTROM and T. WATERS, Swagelok, Cleveland, Ohio


f the analyzer in your analytical system requires gas but your sample is liquid, the only option is to convert the liquid to gas. This process is called vaporization or flash vaporization. The objective is to convert a sample of all liquid to all vapor instantly, without changing the composition. It is not easy to vaporize a sample, nor is it always possible, so make sure it is really necessary and possible before you try. You should always analyze a liquid in a liquid phase unless there are strong reasons for analyzing in a vapor phase. If you proceed with vaporization, it is important to understand the difference between evaporation and vaporization. Evaporation occurs gradually with an increase in temperature. Vaporization occurs instantly with a drop in pressure. It’s not possible to vaporize a sample by increasing temperature. Heat causes evaporation, and adding more heat simply makes evaporation happen faster. In a mixed sample, evaporation will allow some compounds to evaporate before others, resulting in fractionation. Vaporization, done properly, ensures that all of the compounds vaporize at the same time, preserving the sample’s composition. However, it is possible for things to go wrong when vaporizing. Instead of flashing the whole sample into a vapor, you could unintentionally cause a combination of vaporization and evaporation. The result would be fractionation. Once a sample of mixed compounds fractionates, it is no longer suitable for analysis. With fractionation, a common scenario is for lighter molecules to evaporate first and travel on toward the analyzer, while the heavier molecules remain behind in the liquid phase. Even if at some later point in the process a fractionated sample appears to be all gas, the mixture will not be of the same molecular proportions as it was before fractionation. It will no longer accurately represent the product taken from the process line. Let’s take a closer look at the process of vaporization and how we can manipulate the variables (temperature, pressure and flow) to ensure proper vaporization and an accurate analytical result. Understanding vaporization. To vaporize a sample, one typically uses a vaporizing regulator, also called a vaporizer, which is a pressure-reducing regulator with the capacity to transfer heat to the sample at just the right location. Vaporization consists of a three-stage process (Fig. 1). First, the sample enters the vaporizer as a liquid. At this point, the liquid should not be bubbling or boiling.

Second, the liquid passes through the regulating orifice in the vaporizer, resulting in a severe and sudden pressure drop, which vaporizes the liquid. At the same time, heat is applied, which enables the vaporized liquid to remain a vapor. Third, the sample, now a gas, exits the vaporizer and travels to the analyzer to be read. Due to the immediate transition to the vapor phase, the composition of the gas is unchanged from that of the liquid, ensuring an accurate reading. In this delicate process, there are many variables or inputs that determine success or failure. For the purpose of this discussion, let’s say there are two main sets of inputs. The first set of inputs concerns the composition of the sample. Depending on the composition of the sample, it will begin to bubble and finish vaporizing at different pressures and temperatures. We will need to know what these pressures and temperatures are to successfully manage the process. The second set of inputs concerns settings that you control in your sampling system: pressure, temperature and flow. Pressure and temperature are controlled at the vaporizer, while flow is controlled downstream at a rotameter (variable area flowmeter) and needle valve. We set these inputs based on what we know about the first set of inputs. Proper vaporization requires a delicate balance of all inputs. Even when approaching vaporization in a systematic manner like this, the process does require some trial and error, so we will also talk about how to diagnose and address problems. 2) Sudden pressure drop

1) Liquid at high pressure

Vaporizer inlet temperature is below liquid bubblepoint

FIG. 1

3) Vapor at low pressure

Vaporizer outlet temperature is above vapor dewpoint

Drawing showing the three-stage vaporization process.


I 79

PROCESS ENGINEERING Watching out for time delay Time delay is another problem in vaporizing samples. It can be an issue on both the liquid and vapor side of the vaporizer. On the liquid side, the difficulty is caused by the sample’s degree of expansion when it is vaporized. A small amount of liquid creates a large amount of vapor. It’s easy for liquid on the upstream side of the vaporizer to be held up awhile before it is vaporized. If your vaporizer is located near the tap, the best solution is to install a bypass on the liquid side of the vaporizer, so the sample being vaporized is always fresh. In addition, try to minimize

5.0 20% hexane in pentane 4.5

Pressure, bar

4.0 3.5 3.0

Vaporizer body temp.

) ne


Bubble point Dew point Liquid @ 4 bar Temp. drop at vaporization Liquid

s ha o-p w T

oz -g no ( e


Vaporizer core temp.

2.5 2.0 1.5 1.0 50

FIG. 2

Vapor reheat 55



Vapor @ 1.5 bar

70 75 80 Temperature, °C





Phase diagram showing 20% hexane in pentane, with temperature settings.

Understanding your sample. The best way to understand the first set of inputs is with a phase diagram. A phase diagram plots pressure and temperature, showing at any pair of conditions whether a substance will be vapor, liquid or solid. The lines indicate the interfaces between two phases. Phase diagrams for most pure gases are available on the Internet (one example is encyclopedia.airliquide.com). But diagrams for gas mixtures are very difficult to create without commercial software. Fig. 2 represents a phase diagram for 20% hexane in pentane. When the sample is above the bubblepoint (blue line), it’s all liquid. We want the sample to be all liquid when it enters the vaporizer. When the mixture is below the dewpoint (gold line), it’s all vapor. The sample must be all vapor when it leaves the vaporizer. Between the bubblepoint and dewpoint lines is what we call the no-go zone. This zone is the boiling range of the sample. Here, the mixture is in two phases, part liquid and part vapor. Once a sample falls into the no-go zone, it is fractionated and no longer suitable for analysis. The objective in vaporization is to set the temperature, flow and pressure so that the sample skips instantly from the liquid side of the no-go zone to the vapor side of the no-go zone. With pure and nearly pure samples, there is little to no boiling range or no-go zone. The bubblepoint and dewpoint lines are on top of each other or nearly so. Indeed, pure and nearly pure samples will convert to vapor of the same composition, whether through evaporation or vaporization. Some industrial samples approach this level of purity and convert easily. On the other hand, some samples have such a wide boiling range or no-go zone that they cannot be successfully vaporized. There is no way to skip from the liquid side of the no-go zone to the vapor side 80

I JANUARY 2012 HydrocarbonProcessing.com

the volume of the probe and tubing preceding the vaporizer. To address time delay on the vapor side, you may want to increase flow. This may not be the best option. Many samples require low vapor flow rate for proper vaporization. High flow, in combination with insufficient heat at the vaporizer, could result in fractionation, with liquid passing downstream. Such a scenario would ruin the sample for analysis, evident by frost on the tubes downstream of the vaporizer. A better way to reduce time delay on the vapor side is to minimize volume. For example, move the vaporizer closer to the analyzer and/or build a fast loop on the liquid side. HP of the no-go zone. We are unable to manipulate the variables (temperature, flow and pressure) in such a way as to avoid fractionation. Most samples fall between these two extremes. For example, in Fig. 2, the band between bubblepoint and dewpoint is narrow enough that, with the proper settings, we can enable the sample to effectively skip from the liquid side of the no-go zone to the vapor side. At the same time, the band in Fig. 2 is wide enough that we cannot afford to be careless. Indeed, we will need to be skillful in our manipulation of the variables or we will end up with a sample in the no-go zone. Setting temperature, pressure and flow. Let’s continue to work with the sample in Fig. 2 (20% hexane in pentane) and see how we can set our inputs to ensure successful vaporization. In general, at the inlet, we want high pressure and low temperature. At the outlet, we want high temperature and low pressure. But there are limits as to how high and low these parameters can be, and not all of them are under our control completely. Vaporization is basically a balancing act between the variables. Here is a four-step process for setting your inputs: 1. Determine the inlet pressure at your vaporizer. This pressure, which is fixed, is your process pressure, provided your vaporizer is located close to your sample tap. In Fig. 2, that pressure is 4 bar. Higher pressure is better because it allows you to keep the vaporizer temperature higher without boiling the incoming liquid. 2. Set your inlet temperature, or the temperature of your vaporizer. There are two objectives. First, the temperature must be low enough that, when the sample enters the vaporizer, it is entirely a liquid and isn’t bubbling. In Fig. 2, the bubblepoint at 4 bar is 88°C, but we want to build in a cushion, so let’s choose 80°C, a round number far enough away from 88°C to be safe. The second objective is that the temperature must be high enough to contribute to the complete flashing of the sample, ensuring that only vapor leaves the vaporizer. When you vaporize the sample, the temperature drops, in accordance with the laws of energy conservation. The sample temperature must be high enough at the outset so that after the pressure drop, the sample is not in the boiling range or no-go zone. In Fig. 2, the vapor temperature after the pressure drop is 60°C, just on the vapor side of the dewpoint line. 3. Set the outlet pressure at the vaporizer. Your objective is to drop the pressure below the gold dewpoint line. In Fig. 2, the outlet pressure is set to 1.5 bar. If the outlet pressure were any higher in this example, the sample would not vaporize entirely. It would fractionate. 4. Set your flow. Flow is set downstream at a valve and rotameter, not at the vaporizer. In a sampling system, high vapor flow is

PROCESS ENGINEERING desirable because it moves the sample to the analyzer faster. However, high flow can be problematic, too, because with high flow, more heat is required to vaporize the sample. In other words, high flow results in a greater drop in temperature at the time of vaporization. In Fig. 2, the purple line illustrates the temperature drop. As flow increases, the purple line angles more sharply to the left. Another variable influencing the temperature drop is the heat transfer capability of the vaporizer. Some vaporizers are constructed in such a way that heat transfers more efficiently to the sample. When the liquid sample converts to a vapor and its temperature drops, it draws heat from the stainless steel surrounding it. The critical question is how efficiently can the vaporizer replace that heat and keep it flowing to the sample. The more heat the sample can draw, the less its temperature drops during vaporization. In some instances, it is possible for the vaporizer to be hot to the touch on the outside but cold at the core inside. That’s because the vaporized sample is drawing lots of heat and the vaporizer cannot transfer enough heat to keep up. The best solution is to reduce the flow. In sum, the angle of the purple line in Fig. 2 is a product of the flow rate and the heat transfer capability of the vaporizer. With a good vaporizer and low flow, the line will become more vertical. Unfortunately, there is no easy way to calculate the location of the purple line, and it is not generated by any known software program. As a result, vaporization involves some approximation. As a rule of thumb, keep the flow rate as low as possible without causing an unacceptable delay in the sample’s travel time to the analyzer. It’s better to start with a low flow rate and experiment with increasing it than to start with a higher flow rate.

Problem #2 solution. Lower the vaporizer temperature. Know your variables. Vaporizing a liquid sample is challenging. In many sampling systems around the world, vaporizers are fractionating samples and sending unrepresentative samples to the analyzer every minute of every day. You can dramatically increase your chances of success by researching a phase diagram of your system’s particular mixture of compounds. You can further increase your chances of success by understanding what is occurring in the process; specifically, by knowing what the variables are (temperature, pressure and flow) and their role in influencing the process outcome. With this framework in place, you can come very close to the right settings, making adjustments in accordance with the signs and symptoms you observe. HP Doug Nordstrom is market manager for analytical instrumentation for Swagelok, and he focuses his efforts on advancing the company’s involvement in sample handling systems. He previously worked in new product development for Swagelok and earned a number of Swagelok patents for products. Mr. Nordstrom graduated with a BS degree in mechanical engineering from Case Western Reserve University and earned a master’s degree in business administration from Kent State University. Tony Waters has 45 years of experience with process analyzers and their sampling systems. He has worked in engineering and marketing roles for an analyzer manufacturer, an end-user and a systems integrator. He founded three companies to provide specialized analyzer services to the process industries and is an expert in the application of process analyzers in refineries and chemical plants. Mr. Waters is particularly well known for process analyzer training courses that he has presented in many of the countries of Asia, Europe and the Middle East, as well as North and South America. His presentations have equal appeal to engineers and maintenance technicians.

Troubleshooting. Phase diagrams will enable you to approximate temperature, pressure, and flow settings, but some troubleshooting will still be required. One sure indication of a problem is poor repeatability in analyzer results. There are two possibilities when the sample is fractionating instead of vaporizing, with Problem #1 being the more common: Problem #1. Only part of the sample is being vaporized. Liquid is passing through the vaporizer and sitting in the tubing on the downstream side. Eventually, it evaporates. When it does, it draws heat from the surrounding tubing, making the tubing cold to the touch or causing frost or ice to form. Signs of the problem: Vaporizer outlet and downstream tubing is cold to the touch or has frost or ice on it. Note: In many cases, liquid on the downstream side of the vaporizer may pass beyond the area of the vaporizer and into other components, such as flowmeters and filters, where it can cause considerable damage. Problem #1 solution. In the previously discussed approach, your best option would be to reduce the flow rate. Another option would be to lower the vaporizer outlet pressure, if that is possible. A third option would be to increase the heat to the vaporizer, but in this case you risk causing Problem #2. Problem #2. The sample is boiling at the inlet to the vaporizer. It is fractionating before it can be vaporized. Lighter molecules evaporate and create a “vapor wall,” which pushes the liquid back into the process. A portion of that vapor wall then cools and condenses. Finally, the liquid sample moves again toward the vaporizer, where the lighter molecules evaporate, starting the cycle all over again. Meanwhile, the heavier molecules move on toward the analyzer for an inaccurate reading. Signs of the problem: The inlet tube to the vaporizer twitches, sometimes violently, and the measurement values oscillate. Select 162 at www.HydrocarbonProcessing.com/RS

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Case history: Replacement of existing pressure vessel Installing new equipment involves more processes to ensure safety and to meet new codes D. FEARN and J. MCKAY, Fluor Canada Ltd, St. John, New Brunswick, Canada


he replacement of pressure vessels is a common function in an operating refinery, particularly those units that have been in operation for the full design life of the vessel. From the client’s perspective, a vessel may have operated successfully for many years beyond the original design life with no issue. Its replacement should be easily accomplished given the relative success of the original design. In the real world, the replacement of existing refinery vessels in a brownfield environment is seldom replacement in kind. Some minor, yet important, activities should be addressed to ensure project success. In the presented example, work processes used to replace a hydrogen sulfide (H2S) absorber as part of a refinery crude unit that was originally identified as replacement in kind will be discussed. In addition to working to a documented work process, there are many areas where the various design engineers must think outside of the established work practices to ensure the timely, safe and effective installation of new vessels. This article is not meant to replace existing work processes; it will identify unclear areas that exist when replacing equipment in an existing operating unit.

As the vessel operates in an H2S and rich-DEA environment, the refinery performs regular inspections as part of a risk-based inspection (RBI) program as outlined by industry standard practices, refinery specific practices and guidelines established by the American Petroleum Institute. Following an automated ultrasonic testing (AUT) and manual ultrasonic testing (MUT) inspection of the bottom drum, it was found that stress- oriented-hydrogeninduced cracking (SOHIC), resulting in step-wise cracking, and blister formation was present in the drum with concentrations higher in the lower drum region (Fig. 1). Fitness for service calculations resulted in the recommendation to replace the vessel, thus preventing a potential unplanned production interruption. Original vessel. The original vessel was built in 1974 to ASME Section VIII, Division 1, 1971 Ed. The vessel was specified with a joint efficiency of 0.85 (Spot RT) and the material of construction was SA-285 Grade C. Although vessel materials were

The project. This example involves the replacement of an H2S

absorber tower. This tower is commonly found in refinery crude units. In this particular crude unit, the removal of H2S is done by a vertical tower with three integral vessels consisting of two drums and a packed section. The purpose of the tower is to remove H2S, and it is necessary to minimize sulfur dioxide (SO2) emissions from an adjacent atmospheric furnace where the treated offgas is burned. As the atmospheric furnace and adjacent atmospheric column operate at very low pressures, the pressure drop through the H2S absorber must be minimal to avoid excessive backpressure on an adjacent vacuum-seal drum. H2S removal. To do these three activities, the vessel is separated into three major components. The bottom drum contains hydrocarbons, H2S solution and diethylolamine (DEA). The drum removes the hydrocarbons from the H2S, and the DEA assists with this process. The rich-H2S stream is sent from the bottom drum up into an H2S absorber section that is filled with random packing. Treated offgas is sent into the top drum where the untreated DEA is sent back to the absorber section and the gas is forwarded to the adjacent furnace.

FIG. 1

Stress-oriented hydrogen-induced cracking of the H2S absorber column. HYDROCARBON PROCESSING JANUARY 2012

I 83

ENGINEERING AND DESIGN DEVELOPMENTS not post-weld heat treated (PWHT), weld hardness was limited to 200 HB. The bottom drum is 4 ft in diameter and 40 ft in height (including the 14-ft skirt). The top drum and packed section is 70 ft long and 2 ft in diameter. The H2S absorber internals consisted of four hold-down plates and 120 ft3 of random packing. To distribute solution to the H2S absorber, tangential nozzles were used on the inlet nozzles where the solution would collect on the spray header that is then gravity-fed down through the packing. To limit direct contact (and subsequent erosion) of DEA and H2S on carbon steel, the tangential nozzles directed flow toward a 304 stainless steel (SS) clad plate that was welded to the vessel internal diameter. Theoretical design considerations and scope. Ini-

tially, the project requirement was the replacement in kind of the vessel; other equipment directly attached or adjacent to the vessel would also need to be examined. For example, the piping system that also processes rich DEA and H2S was potentially at risk for associated metallurgical damage mechanisms. The foundation and corresponding anchor bolts needed to be reviewed to determine if they were acceptable for continued service for the estimated design life of the replacement vessel. Electrical and control systems were also reviewed to determine if existing systems are code compliant and adhere to current refinery practices. While other disciplines face challenges similar to the vessel designer, this

FIG. 2

FIG. 3


Guide locations from the adjacent atmospheric column.

Trail-fit ladders and platforms manufactured by the vessel fabricator.

I JANUARY 2012 HydrocarbonProcessing.com

article will focus solely on the vessel replacement. However, noted items are considered inter-discipline related. For the vessel designer, the scope to replace a vessel includes far more than the replacement in kind of an existing asset and ensuring that the new asset will meet the latest codes. The designer must engage operations to ensure that manway size and location, ladder and platform access, packing access, etc., are acceptable to the current and future needs. Some needs may not be identified until the piping, electrical and controls designers also do their respective design activities. As the request from the client was to replace the vessel in kind, the existing vessel was modeled into an available simulation model to determine if revisions to ASME Section VIII between 1971 and 2008 would result in an overall design change to the new vessel. For this vessel, particular attention was paid to the internal head design for the bottom/top drum assemblies, as well as the transition (48 in. to 24 in.) between the bottom drum and random packing section. Where the new vessel was to be constructed of SA-516-70N, the greater allowable stress compared to the original SA-283 offsets any code changes that would otherwise increase the overall thickness and potentially impact the total dimensions. An important activity of the vessel designer is to visually verify and place hands on every item of the vessel and to check its accuracy against the original design drawings. This includes additional vessel penetrations, platform loads or equipment that was not part of the original design. Depending on the level of documentation control within the existing facility, it is possible that the original drawings do not exist or are of such poor quality that new drawings must be drafted. Whereas construction is not typically engaged until further in the fabrication process, brownfield development should include design considerations recommended by the construction team and lift contractor. Items addressed include the timing of internals installation, adding lift lugs or relocating platform clip locations to facilitate installation where space is limited. The output from this process should include a defined and inter-discipline reviewed datasheet and general arrangement drawing (as-built or new) that will be issued to the vessel fabricator. This allows all disciplines to review potential interferences between nozzles, clips, guides, supports, girth flanges, etc. Actual design conditions and scope development.

The original project scope basis was for a replacement in kind vessel. However, during project development, many changes were made. To identify the changes required for this vessel, each discipline input was identified separately. Mechanical related changes. Given the presence of H2S and SOHIC, the base materials were upgraded to SA-516-70N – HIC resistant carbon steel and included for PWHT in accordance with the recommendations of NACE MR0175. The radiographic testing (RT) was increased to full 100% RT while maintaining the 200 HB harness limit. To prevent solution entrapment between the original internal shell, it was decided to use 304L weld overlay on top of 309L. Minimum weld overlay thickness was specified to ensure adequate thickness for long-term protection. Additional NDE was specified for the overlay, such as LPI, UT (for disbondment) and ferrite testing. This vessel is tall relative to the base diameter, and, without supplementary support, it requires additional material on the base ring, anchor bolts and girth flanges to resist buckling due to the wind and seismic overturning moment. The original vessel design included

ENGINEERING AND DESIGN DEVELOPMENTS guides, as shown in Fig. 2, at three upper elevations whereby the adjacent column provided support. The inclusion of the guides permitted the redistribution of wind load (overturning moment on the vessel base) and the acceptance of the existing anchor bolts and vessel thickness. Per requirements by the client and construction team, hand holes were installed to facilitate inspection, removal and installation of the random packing. Lift trunions were also installed on the base of the packed section to facilitate installation. Several new pipe guides and supports were also installed on the vessel to remove loads from platforms. Following a review of the loads placed on the ladders and platforms, in addition to occupational health and safety changes since original construction, the ladders and platforms were redesigned. Electrical related changes. As the refinery specifications have changed over the years, the requirement was made to include for cable tray clips on the vessel to facilitate electrical cable installation and pre-dressing prior to lift.

FIG. 4

CAD drawing for the lift contractor to confirm lift plan relative to available working space and local obstructions.

FIG. 5

Engineered lift lug and space limitations with pre-installed platform and adjacent column.

Vessel fabrication. For this particular project, the vessel fab-

ricator was provided with the original as-built vessel drawing and revised datasheets. The fabricator was required to produce new drawings incorporating all of the changes. For the vessel designer, this requires attention to detail to ensure that while overall dimensions are consistent with the original design, all changes have been incorporated into the new design. To help facilitate quality concerns between the fabricator and client, a third-party inspector was enlisted throughout the fabrication process. The scope of the third-party inspector was to ensure the agreed to inspection and test plan was being adhered to, as well as to be a client representative for any hold points during the fabrication or final assembly and test process. As new platforms were specified for the replacement vessel, it was decided that a shop-trial fit test should be done. This ensured that the platforms would fit during installation and prevented costly rework onsite that might, otherwise, have to be performed within the turnaround window. Demolition and installation of vessel. Depending on

the time available, space considerations and resource availability, the construction team may choose to pre-install as many vessel related components as possible to reduce construction costs and to prevent doing work within an operating unit. This may include pre-dressing fireproofing, process pipes, heat tracing, insulation, valves, platforms, internals, instrumentation and cable trays. Where this vessel was being removed and installed in an operating environment, these processes were followed: • Lift contractor reviewed available space and determined maximum allowable lift capacity (Fig. 4). • Construction determined the recommended extent of predressing. • Mechanical engineering determined the total weight and center of gravity for the vessel, complete with all pre-dressed components. Depending on the amount of materials pre-installed, and the level of certainty of equipment/weight estimates, a lift factor was incorporated into the overall maximum lift weight. A factor of 10%–30% is not uncommon to include for errors in drawings, fabrication tolerances, etc. This weight becomes very important as space availability may limit the crane type and capacity. Improper weight estimates to the lift contractor may result in too small (or too large) of a crane.

• The lift contractor verified the weight prior to the demolition/installation to confirm that crane capacity would not be exceeded. Elevations, offsets and grouting of baseplate. During the design and construction process, the elevation of the foundation and underside of the vessel base ring was surveyed. This ensured that the vessel would rest at a similar elevation to the original asset. With the original vessel removed, the foundation was prepared to accept the new vessel and associated grout. Once the vessel was installed, a survey was completed at the underside of the base ring at the shim locations as well as all girth flanges on the vessel. This served as a check to ensure there was no “wobble” in the vessel sections. Once the vessel had achieved proper alignment, the vessel was grouted to the foundation with the remaining components (ladders, internals) installed that could not be pre-dressed. HP Dan Fearn, P.Eng., is a design engineer with Fluor Canada Ltd. He holds a BS degree in mechanical engineering and is a registered Professional Engineer. Mr. Fearn has more than 10 years in mechanical engineering; his expertise lies in the in the specification and selection of mechanical equipment and in the development and implementation of maintenance programs with a focus on site support and installation.

Jeff McKay, P.Eng., is a senior design engineer with Fluor Canada Ltd. He holds a BS degree in mechanical engineering and is a registered Professional Engineer. Mr. McKay has 14 years of experience with Fluor Canada Ltd., and, at present, is the lead mechanical engineer at a client jobsite. HYDROCARBON PROCESSING JANUARY 2012

I 85


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2011 REFINING PROCESSES HANDBOOK www.HydrocarbonProcessing.com

Produced by the staff of Hydrocarbon Processing magazine, this comprehensive industry reference source contains flow diagrams and descriptions of more than 130 leading-edge, licensed refining technologies. Specific processing operations include coking, hydrotreating, hydrocracking, fluid catalytic cracking, resid catalytic cracking, alkylation, catalytic reforming, ethers and more. Forty licensors contributed process flow diagrams, products, process descriptions, economics, installations and other vital information. The flow diagrams and summary descriptions define typical licensed processes used by modern refineries.

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FREE Product and Service Information—JANUARY 2012 HOW TO USE THE INDEX: The FIRST NUMBER after the company name is the page on which an advertisement appears. The SECOND NUMBER, appearing in parentheses, after the company name, is the READER SERVICE NUMBER. There are several ways readers can obtain information: 1. The quickest way to request information from an advertiser or about an editorial item is to go to www.HydrocarbonProcessing.com/RS. If you follow the instructions on the screen your request will be forwarded for immediate action. 2. Go online to the advertiser's Website listed below. 3. Circle the Reader Service Number below and fax this page to +1 (416) 620-9790. Include your name, company, complete address, phone number, fax number and e-mail address, and check the box on the right for your division of industry and job title. Name ________________________________________________________

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Case 66: Fiberglass mixing tank flexing vibration The design of fiberglass and polymer tanks had many problems in the 1980s. Most of these problems were resolved with experience, new materials and manufacturing techniques, along with new national construction standards. Such tanks are now considered quite reliable.

Tank wall Baffles Impeller Favg

Problems. However, operating problems do continue to pop up

with fiberglass tanks, as was the case when the following analysis was done. In this application, a prototype fiberglass tank was designed to mix caustics in a small pilot plant. The mixer was mounted on a separate structure and was not in contact with the tank. This tank vibrated at a low frequency and developed stress cracks within the vicinity of the mixer blades at the baffles. The presented model was constructed to help explain the causes for the stress cracks and to identify possible modifications. This simple analysis answered many questions. Fig. 1 is a view looking down into the tank, with the impeller and baffles shown. Also shown is an imagined water slab captured by one sector of the impeller. The impeller was near the floor of the tank; so, the flow was mostly tangentially outward, as shown in Fig. 1 and replenished from the impeller top. The water is considered to be a slab the thickness of the blade width, h, and a sector of the circle, Dimp . The mixer is rotating at a given rpm with a tip velocity of V. There is a velocity profile, but, for simplicity, the slab is assumed to move at a tip speed, V. The trajectory of the slab is shown by the dashed arrow and impacts the tank wall with a force Favg (Fig. 1). This occurs because the water must be displaced. Since it does not go down or up within the tank, it is assumed to travel tangentially, where it contacts the wall and baffle and then swirls up. Tank model. The analysis to determine Favg will assume that

the weight of the water slab with h, the blade width, n, the number of blades; and , the fluid density, is: W = 0.785  Dimp 2  h    (1/n) lb where: m = W / 386 lb-sec2/in. The tip velocity is: V = π  Dimp  rpm/60 in./sec The time for the slab to deform and decelerate when it hits the wall is assumed to be the time for the rotor to make 1/n turn. Result: The slab has exited and is filled again:




FIG. 1

Water slab

Mixer and tank model.

 = 0.036 lb/in.3 Dimp = 84 in. h = 4 in. rpm = 45 n = 4. The results is Favg = 307 lb, and it is a cyclic force on the wall at a baffle of 180 vpm, since there are four blades passing a baffle at 45 rpm. Some portion of this load impacts the baffle that may have been responsible for the flexing, cracking and vibration issues. Because this was a pilot plant, it was only to operate for three years. The mixer speed was lowered to 30 rpm with a gear change, which reduced the force to 136 lb and had no adverse affect on the mixing. It did, however, eliminate much of the flexing. This allowed continued operation until the unit was retired from service. Even with this rather simplistic model, the internal loads causing the vibration and flexing were better understood. Sometimes. this is all that is required from an analytical model. The conclusions and recommendations were clear and were presented to management in a quantitative form, which is always better than speculation. HP

Δt = (π  Dimp /V )  (1/n) sec Favg = m  a = W  V / (386  Δt ) lb. For this case the fluid is water 90

I JANUARY 2012 HydrocarbonProcessing.com

The author, PE, was the worldwide lead mechanical engineer for ExxonMobil before his retirement. Information on his books, seminars and consulting, as well as comments to this article, are available at http://mechanicalengineeringhelp.com.


Unconventional Feedstocks and Heavy Oil Conversions You are invited to submit an abstract for Hydrocarbon Processing’s third annual International Refining and Petrochemical Conference (IRPC) that will be held 12–14 June 2012 in Milan, Italy. IRPC is a leading-edge technical conference, providing an elite forum within which industry leaders can share knowledge and ideas relating to the international refining and petrochemical industries. The conference emphasizes the latest technological and operational advances from both a local and global perspective, and is attended by project engineers, process engineers and hydrocarbon processing industry (HPI) management officials from around the world. With changes in crude supply around the world, refiners and technology companies will be able to present the latest developments in refining technologies. The topics to be covered at IRPC 2012 include (but are not limited to): • Heavy oil conversion/bottom-of-the-barrel • Plant and refinery sustainability • Profitability • Energy policy • Middle distillate developments • MARPOL regulations • Shift in gasoline to diesel ratio (subject to specific countries) • Renewables/biofuels • Future of fuel oil • Clean fuels • Plant safety • Flare systems • Gas treatment technologies • Rotating equipment

• Refinery and petrochemical integration • Bio-based petrochemicals/chemicals • Alternative feedstocks—shale gas, GTLs, CTLs, etc. • Catalysts—rare earth issue and new developments • Metallurgy • New Materials • Mechanical equipment • Energy efficiency • Maintenance and reliability • Effluence management (water, air, solid waste) • Carbon management • Process control applications/automation

Abstracts should be approximately 250 words in length and should include all authors, affiliations, pertinent contact information, and the proposed speaker (who will present the paper). Please submit via email to [email protected] by 3 February 2012. The conference advisory board will review all submitted abstracts and select which will be presented at the conference in June. For more information on the conference and to learn about other ways to get involved, please visit www.HPIRPC.com.

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