Novel contacting technology selectively removes H2S

May 13, 2002
A new compact selective H2S removal process is useful in a number of gas processing applications. The process is based on the cocurrent FramoPure gas-liquid contactor, from Framo Purification AS, Bergen, with commercial nonproprietary solvents.

A new compact selective H2S removal process is useful in a number of gas processing applications. The process is based on the cocurrent FramoPure gas-liquid contactor, from Framo Purification AS, Bergen, with commercial nonproprietary solvents.

The cocurrent technology's absorption section takes up less space and weighs less than other processes. This is particularly important for high-pressure applications. The process also has a lower solvent circulation rate compared to conventional technologies. The installation, therefore, weighs considerably less.

The most common process for offshore removal of H2S is direct-injection H2S scavenging. The process is simple and robust, and capital expenses are low, but operating expenses are high. There are growing concerns, however, involving the environmental impact of H2S scavenging.

We developed the technology for selective H2S removal based on a compact alkanolamine plant (CAP) in connection with the Kristin gas condensate field development in the Norwegian Sea. The principal solvents were aqueous solutions of methyldiethanolamine (MDEA).

The process can remove H2S to a pipeline specification with residence times of 50 ms. H2S loading capacity increases due to the high selectivity for H2S in the cocurrent contactor. Required H2S selectivity is, therefore, primarily facilitated through the contacting device and only partially depends on the alkanolamine. A lower solvent circulation rate is possible with the higher H2S loading capacity.

We compared circulation rate, weight, and cost for three cases from existing gas producing fields or fields under development on the Norwegian continental shelf. For certain applications, the dry weight of CAP is 33% of a conventional process. For offshore installations, weight, space, and cost may limit use of conventional amine sweetening processes based on counter-current absorption.

The life-cycle costs (LCC) for CAP are competitive with H2S scavenging for low H2S concentrations. The CAP is a viable alternative to H2S scavenging given the current concerns regarding environmental discharges to the sea.

Zero emissions

Sour reservoirs and hydrocarbon-producing reservoirs boosted with water on the Norwegian continental shelf typically have low concentrations of H2S-from 10 to 60 ppmv. Despite these low H2S concentrations, the natural gas product must satisfy the sales specification (2-4 ppmv) and avoid corrosion in pipelines and process facilities.

Due to low capital cost, offshore treatment facilities for natural gas with low H2S concentrations, low gas flow rates, or low life cycle times are typically H2S scavenging with irreversibly reacting chemicals. H2S scavenging uses injecting chemicals that selectively and irreversibly react with H2S.

Since aqueous by-products can discharge to the sea with produced water, scavenging can have a major environmental impact when large gas streams are treated. For several fields, H2S scavenging contributes considerably to the environmental impact factor.1

For the Norwegian continental shelf, oil and gas operators and the Norwegian Pollution Control Authority targeted 2005 as the year for "zero emissions" of harmful discharges to the sea. This creates an incentive for operators to implement more environmentally friendly H2S-removal technologies.

Alkanolamine sweetening processes based on conventional countercurrent absorption, common for onshore plants, provide an alternative technology in some cases. For offshore installations, however, weight, space, and capital cost limit the use of these large amine processes.2

Offshore H2S removal

Conventional H2S removal technologies include scrubbing processes with regenerative amine or physical solvents, direct liquid oxidation, catalytic conversion, chemical or physical adsorption, and nonregenerative H2S scavenging with liquid solvents or metal oxides.

Difficulties arise when a small amount of H2S must be removed from gas containing CO2 to satisfy pipeline specifications. Several processes that selectively remove H2S in the presence of CO2 may be uneconomical for H2S polishing applications. These processes are complex and expensive. Operators, therefore, use low capital cost H2S scavenging.

In recent years, technologies for gas sweetening with amines have become more sophisticated, with major advancements in more selective solvents.

Kinetic selectivity results because the reaction of amines with H2S and CO2 occur at vastly different rates. Whereas the reaction with H2S involves a proton transfer and is fast, CO2 reacts at a limited rate with amines and other alkaline substances.

Industry extensively uses the tertiary amine, MDEA, for selective absorption of H2S from natural gas, refinery gases, and coal gasification. Specially developed solvents, such as sterically hindered amines, can further lower the CO2 reaction rate.

Amine sweetening processes using proprietary, selective amine solvents combined with countercurrent absorption columns are an alternative technology. The recent development in the Norwegian Sea, Åsgard B, is an application in which the operator obtained a low circulation rate and a low plant structure weight on the platform.3

More-selective solvents, rather than a methodology for contacting gas and liquid, have led to major advances in selectivity. Although conventional absorption towers have undergone major developments internally, gas-liquid contacting times are still high, which allows the inevitable coabsorption of CO2. Improved control of selectivity requires an additional degree of freedom: the contacting time between gas and liquid.

Novel technology

Framo's pilot plant at the Statoil Mongstad refinery demonstrated the new technology's selectivity. Only 10% of the CO2 in the feed gas was coabsorbed in the FramoPure cocurrent contactor vs. 60% in an existing countercurrent column.4 We established the new H2S removal process for the Kristin field. The circulation rate and weight is 65-75% lower for the CAP vs. conventional alkanolamine processes.2

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The FramoPure contactor (Fig. 1) is a gas-flow driven, "one shot" contactor. A local increase in the dynamic pressure of the gas flow transforms the liquid (solvent) to the contactor into small droplets.

Any mixing creates a permanent pressure drop. With the new contactor, the pressure drop is associated with momentum transfer from the gas to liquid. This happens within a short distance, as opposed to momentum transfer between phases in relatively long flow conduits with high wall shear stress. The process, therefore, achieves a high degree of mixing with a correspondingly large interfacial gas-liquid surface at a low to moderate pressure drop.

The droplet generation mechanism is a process in which an annular liquid filament of low surface tension force is:

  • Initially generated by the liquid injection chamber.
  • Exposed to gas flow with high inertial force.

The high ratio between the inertial force and surface tension force breaks up the liquid filaments into small droplets. Important characteristics of the FramoPure contactor are:

  • The formation of small liquid droplets.
  • High gas-phase turbulence that improves the gas-side mass transfer.

The cocurrent contactor can be smaller because it operates at higher gas flow rates than a countercurrent contactor. Operational difficulties due to flooding or nonhomogeneous gas-liquid distribution, common in countercurrent absorption towers, are not a problem with cocurrent contacting.

Test setup

We conducted contactor testing at Framo's test facility in conjunction with Statoil's Kristin gas condensate field development. H2S and CO2 feed concentrations were 20 ppmv and 10 mole %, respectively.

We performed the tests in a 3-in. contactor at a nominal pressure of 13 barg. The gas flow rate was 250 cu m/hr, corresponding to a gas velocity of 13 m/sec, and gas-liquid mixture temperature was 35° C.

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Fig. 2 shows the test setup and sampling arrangement for the contactor and downstream contactor pipeline. Sampling points occur before and after the Framo Pure contactor and along the contactor pipeline. This helps determine absorption kinetics of the gas-liquid reaction.

We determined H2S-removal performance using batch testing because the contactor pipe quickly attains steady-state conditions, both hydrodynamically and in terms of solvent loading.

We initially enriched the air with CO2 gas and recirculated it until a stable and homogeneous CO2 concentration resulted. We then fed an N2-H2S gas mixture into the circulating gas upstream of the contactor at a constant rate corresponding to a prescribed H2S concentration.

Finally, we injected the solvent into the contactor at a constant rate for 60 sec. About 30 sec after steady-state conditions were reached, the data acquisition system remotely triggered the gas sampling.

We set the sampling schedule according to the propagation time of gas through the pipeline to the individual sampling points. The sampling probe in the two phase gas-solvent flow avoided any solvent droplets entrained in the sampled gas.

We regularly flushed the sampling arrangement tubing, which occupied a minimal interior gas volume. We also regularly flushed and saturated the tubing with H2S-rich gas to reduce the effects of H2S adsorption on the internal metal surfaces.

We measured and compensated for the effect of H2S adsorption and absorption in each individual sampling arrangement by extracting a sample from the circulating H2S-enriched gas flow before and after each gas-solvent flow experiment.

We ran the experiments with a constant gas flow rate. The solvent feed rate varied for each experiment. This way, we could derive the H2S-removal and scaling property of the specific rate of reaction as a function of the gas-solvent flow rate ratio. We analyzed the gas concentrations with gas chromatography.

We tested three different nonproprietary solvents-DEA, DIPA, and MDEA-with molarities of 2 mole and 4 mole.5 6 All solvents were free of H2S and CO2 before injection.

Kristin test results

Fig. 3 shows H2S removal vs. gas-liquid volumetric flow rate ratio (GLR) results for the tested solvents. Apart from three experiments with a 4-molar MDEA solution, all the results refer to nominal CO2 concentrations of 10%. To compare the experiments representing different H2S feed concentrations, we chose H2S-removal efficiency relative to inlet concentration.

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Results demonstrate that MDEA is the best performer. Of the two solvent concentrations tested, 4 molar gave the best performance.

For the test conditions, 4-molar MDEA will yield 90% H2S removal at a GLR of 210. Reducing the CO2 concentration from 10% to 5% further increased the limiting GLR. The CO2 coabsorption reflects the difference in solvent performance. Whereas DIPA absorbed a total of 10-12% of the CO2 gas, the corresponding value for MDEA is less than 5%.

The FramoPure contactor is crucial for selectivity and the resulting performance. The selectivity towards H2S is typically 35-100 times higher for the contactor mixing section than the downstream pipeline. Fig. 4 shows a H2S-removal profile vs. residence time downstream of the contactor.

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The first measurement point corresponds to a residence time of about 50 ms. We calculated "data compensation" (Fig. 4) by correcting for the gas absorption and adsorption occurring in the gas sampling arrangements.

The accuracy of the H2S measurements from the GC was ± 0.7 ppmv. Finally, the permanent pressure drop in the FramoPure contactor was less than 0.14 bar for all of the experiments.

Selective H2S absorption

Selective H2S removal commonly uses an appropriate chemical with selective characteristics in conventional absorption columns. The FramoPure cocurrent contactor is more selective towards H2S than countercurrent contactors due to the high gas-liquid interfacial area available for mass transfer and the short reaction time.

The cocurrent contactor has two principal parts with respect to sour-gas absorption rate and selectivity:

    The injection and contactor mixing section, which facilitates droplet breakup.
  • Downstream contactor pipeline in which the gas and droplet flow gradually approach a steady-state flow pattern for actual flow conditions.

For a given application, one can vary the length of the contactor pipe and solvent feed rate to arrive at a specified H2S gas concentration while minimizing the amount of coabsorbed CO2.

Droplets generated in the contactor mixing section undergo phases of coalescence and deposition on a liquid film that exists on the downstream pipe wall. The effective reaction area will, therefore, inevitably decrease downstream if no efficient remixing is imposed. This effect reduces the actual absorption rate and H2S selectivity in the contactor pipeline.

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One can consequently adapt the process-piping configuration to the requirements of a given application; e.g., a short downstream piping layout is needed for a selective H2S removal application such as Kristin. Multiple solvent-feed configurations and gas-liquid remixing can achieve optimal performance and process control.

A single-stage cocurrent contacting process yields equivalent or lower maximum acid-gas loading compared to a countercurrent process for complete acid-gas removal or sweetening of gas containing only one acid-gas component. High CO2 coabsorption in a countercurrent column significantly reduces the maximum H2S loading capacity, however, if the process must remove only a small amount of H2S from a CO2-rich gas.

Due to the high selectivity for H2S in the FramoPure contactor, its loading capacity for H2S will increase. A higher H2S loading capacity requires a correspondingly lower solvent circulation rate.

Application studies

A thorough understanding and modeling of the reaction kinetics helps one interpret test results of cocurrent contacting. Commercially available simulation codes provide an inadequate description of reaction kinetics for amine sweetening in cocurrent contactors. In fact, predictions for H2S and CO2 absorption rates are only reliable for a limited number of conventional alkanolamine solvents in countercurrent column simulations.

Developing a model with specific kinetic reaction rates for combinations of solvent and contactor, partial pressures, and loading capacities takes a significant effort. We combined our model with the thermodynamics and numerical solver in the process simulator, HYSYS.Process. The model, combined with experimentally determined reaction rates and published solvent vapor-liquid equilibrium data,6-9 is a powerful design tool.

For a given feed gas and solvent, the simulation setup and kinetic model determine the solvent-circulation rate for a fixed H2S specification. Solvent-circulation rate is the fundamental parameter when we estimate the dimensions and capacities required for the process equipment (gas-solvent separator, regenerator, heat exchangers, pumps, reboiler, etc.).

Process description

Fig. 5 shows a principal flow scheme for CAP. The process is based on a conventional absorption-regeneration cycle. The main difference is the absorption section, which consists of two serially configured FramoPure contactors with multiple solvent feeds for optimal process control and turndown.

A conventional scrubber or modified glycol contactor scrubber separates the solvent from the gas. The glycol scrubber allows a further reduction in process weight and plot space-particularly important for offshore installations. The scrubber section design is based on conventional sizing with appropriate safety margins. Amine carryover can seriously impair dehydration system performance and increase operating costs due to solvent losses.

A superheater should be used to reheat feed gas from the high-pressure scrubber to minimize hydrocarbon condensation and absorption during acid gas absorption. A gas temperature of 28-30° C. is sufficient for H2S absorption and downstream dehydration.

The regeneration cycle uses hydrocarbon flash (optional), lean-rich solvent heat exchange, and thermal steam stripping in a conventional refluxed stripping column. The stripping column has one or two sections of structured packing and a short reflux section. Structured packing is optimal for floating facilities and also provides a compact design suitable for offshore installations. Steam is generated from the aqueous solvent and accumulated water in the stripping section.

The solvent reboiler is a conventional kettle reboiler or a more-compact horizontal or vertical thermosyphon, which is favorable offshore.

Overhead gas is cooled in a compact condensing heat exchanger, which also separates gas and liquid. Condensed water is recycled to the reflux section where it reduces solvent losses. Acid gas from the condenser overhead, which is CO2 and H2S saturated with water at 30-40° C., also contains some hydrocarbons, particularly if there is no flash drum. Gas is preferably routed to a flare system or an H2S incineration system.

Case studies

All three cases are from existing gas fields or fields under development. The H2S specification is 2 ppmv except for Case 3, which also considered 10 ppmv. None of the fields requires CO2 removal because gas blending controls the sales-gas specification. We used an aqueous solution of 4 molar MDEA (46 wt %).

The Kristin field in the Norwegian Sea is currently under development. The proposed gas production is 615 MMscfd with 11-20 ppmv H2S and 3.5-5 vol % CO2.

We considered several gas-sweetening alternatives, including amine wash, scavenging, and dry adsorption (land-based treatment alternatives).

Statoil's Åsgard B semisubmersible gas processing platform, started up in 2000, is designed to process 1,300 MMscfd. The platform is equipped with a state-of-the-art hindered amine plant for selective H2S removal. The amine plant capacity is 850 MMscfd of sour gas containing 50 ppmv H2S and 5 vol % CO2. Statoil installed a direct injection H2S scavenging system for backup and peak-shaving service (450 MMscfd).

Case 3 is an existing North Sea facility, which currently boosts reservoir pressure using seawater injection. The operator is currently evaluating the consequences of future increases in H2S concentrations. We assumed a feed H2S concentration of 60 ppmv for this application. The corresponding purification H2S specification is 3-10 ppmv.

Table 1 shows process data used as the basis for evaluating the selected gas-sweetening technologies. We calculated characteristic parameters and established weight and cost estimates.

Circulation rate

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Amine plant cost, weight, and energy consumption closely correlate to the amine circulation rate. Based on the cases in Table 1, we calculated amine circulation rates for CAP using our kinetic model. A commercial process simulator provided circulation rates for conventional gas sweetening with MDEA.

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Table 2 shows that the cocurrent technology has a lower circulation rate than conventional technology when the gas contains small amounts of H2S in the presence of CO2. The Kristin cases clearly illustrate this.

As the H2S concentration increases, the equilibrium limitation for cocurrent contacting becomes more significant, but the technology can still compete with conventional technology. In Case 3, the sour gas CO2 concentration is 2.2 vol %, lower than for the Kristin and Åsgard B cases.

With respect to circulation rate, conventional countercurrent columns are more competitive as CO2 concentration decreases. For Case 3, however, the calculations show that cocurrent technology is the best alternative because the contactor's weight is significantly lower than for a countercurrent column.

The Åsgard B case illustrates that CAP can compete with state-of-the-art hindered amine solvents.

The circulation rate is in the same range, the process equipment weight is lower, and the cost of commodity solvents used for CAP is significantly lower, compared with that of proprietary hindered amines.

Weight estimates

An amine natural gas sweetening plant consists of several principal units:

  • Absorption equipment (column or contactor).
  • Regeneration system.
  • Upstream and downstream gas conditioning equipment necessary for the amine unit.
  • Acid-gas handling equipment (H2S incinerator).
  • Miscellaneous utility systems.

Weight estimates relate to the absorption section and regeneration system. We assume the associated weight and cost for auxiliary equipment3-5 is the same for the cases included in the comparison. The weight of this equipment is constant or proportional to the amine plant weight.

An amine plant's weight depends on the amine circulation rate and size of the gas-liquid contactor. For a high-pressure gas sweetening plant, a conventional absorption column can contribute to half of the dry process equipment weight.

For the compact cocurrent process, the contactor is not much heavier than a pipe spool piece, and the dehydration column or compressor suction scrubbers can integrate the gas-liquid separator, thus adding a few extra meters of shell weight. In the worst case, the operator must use a dedicated scrubber, but the associated weight is less than half the weight of an absorption column.

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Figs. 6 and 7 present a process equipment weight comparison. The dry process weight of CAP is 33% of the dry weight for a conventional MDEA process. Additional weight of the conventional countercurrent amine process is due to the absorption column, which has a dry weight of approximately 100 tons, and the regeneration system with 2.5 times the capacity of CAP.

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For the Åsgard B case (Fig. 7), the dry weight for CAP is also 33% of a conventional MDEA process and 43% less than the existing hindered amine process. Although the hindered amine process has almost the same amine circulation rate as CAP, the absorption column (125 tons) significantly increases equipment weight. Removing the absorption column and including an additional 4.5 m of extra shell length in the glycol column corresponds to the 43% weight reduction.

Lifecycle costs

It is difficult to attain realistic lifecycle cost (LCC) calculations; they are not easy to use because each field is unique and has its own special requirements and cost factors. Contractors and oil companies usually apply their own methodologies when calculating LCC.

An amine system has critical implications for the fluid processing scheme, platform auxiliary systems, and platform layout and construction. The total installed weight associated with an amine plant (inclusive additional platform deck structure, etc.) is usually 2-4 times the wet process equipment weight.

Direct injection H2S scavenging is simple and robust, and most plants have low capital costs. Many offshore H2S scavenging systems are poorly designed, however, because good performance prediction tools are not commercially available.

Engineers tend to consider the scavenging process as one with almost no design requirements. A well-designed direct injection H2S scavenging system can, in fact, achieve specific chemical consumption values of 8-12 l./kg of H2S removed,10 although many report significantly higher numbers.

We calculated LCC cost for H2S scavenging based on 16 l./kg of H2S removed, which is realistic for the fields included in this work.

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We estimated LCC for CAP and H2S scavenging for two alternative installations on a simple basis. The calculations include operating costs and overall capital costs for installation on semisubmersible gas condensate processing platforms. The calculations are valid for constructing, installing, and operating in the North Sea.

The costs for CAP may be lower in other regions of the world. Figs. 8 and 9 show normalized results, given an LCC for H2S scavenging of 100%.

Fig. 8 shows that the cost for CAP is almost identical to that of H2S scavenging. The environmental impact from discharged H2S scavenger can be considerable; this impact is lower with an amine process such as CAP. CAP is, therefore, the best alternative for this particular field.

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Fig. 9 shows that the LCC for CAP at Åsgard B is 58% lower than H2S scavenging. CAP is also the best approach for Åsgard B.

Fig. 8 also shows that if the incineration plant can be eliminated, the LCC for CAP is even lower. Flaring H2S-rich gas yields SO2 emissions, which one should carefully evaluate.

For smaller fields and fields with low H2S concentrations, however, flaring the stripper overhead acid gas is acceptable, particularly if H2S scavenging is the only alternative.

Operating expenses for H2S scavenging are nearly proportional to the amount of H2S removed. The two examples, therefore, illustrate the general trend higher LCC for H2S scavenging as H2S concentration increases.

The amount of gas processed also has an important influence on the optimal H2S removal route. For small fields, H2S scavenging is the best choice. CAP, however, shifts the limiting gas flow rate and H2S concentration for regenerative H2S removal towards smaller fields and lower H2S concentrations.

Technology qualification

Statoil, TotalFinaElf SA, and Phillips Petroleum Co. Norway have started a joint industry project to qualify the technology. Gas-liquid separation optimization testing and absorption performance testing is in progress at Framo Purification's test facility. The dense gas, SF6, will simulate high-pressure conditions.

Qualification of cyclic performance, turndown and robustness for different feed conditions will continue either at Statoil Mongstad or another appropriate test site in Europe. Åsgard B has already demonstrated scale up of the FramoPure contacting technology.

Statoil installed two 24-in. FramoPure contactors in series as a part of the H2S scavenging system.

The two separate units have independent solvent distribution, which facilitates process control and optimization, similar to the concept developed with the Kristin case study.2

For the initial start-up period, the H2S concentration was low (7 ppmv). Such "polishing" applications normally require higher H2S scavenger consumption. Injection in the only upstream quill initially proved unable to reduce the H2S level notably, even with high levels of injected scavenger.

With the FramoPure contactors, the H2S level dropped from 7 to 0.1 ppmv with more than a 35% reduction in scavenger liquid compared to conventional direct injection methods.10

References

  1. Knudsen, B., Tjelle, S., and Linga, H., "A new approach towards environmentally friendly desulphurisation," presented to The 12th International Oil Field Chemistry Symposium, Geilo, Norway, Apr. 1-4, 2001.
  2. Nilsen, F.P., Lidal, H., Linga, H., "Selective H2S removal in 50 milliseconds," presented to the GPA Europe Annual Conference, Amsterdam, Sept. 26-28, 2001.
  3. Van Son, K.J., Chludzinski, G., Charles, P.R., and Lidal, H., "Åsgard B process selection for hydrogen sulfide removal and disposal," World Oil, September 1999.
  4. Linga, H., Nilsen, E., Nilsen, F.P., Soerum, P.A, Johansen, S., and Pedersen, T., "New selective H2S removal process for the refining industry," presented to the NPRA 2001 Annual Meeting, New Orleans, Mar. 18-21, 2001.
  5. Astarita, G., Savage, D.W., and Bisio, A., "Gas treating with chemical solvents," John Wiley & Sons, New York 1978.
  6. Kohl, A.L., and Nielsen, R.B., "Gas Purification," Fifth Edition, Gulf Publishing Co., Houston, 1997.
  7. Jou, F.Y., Lal, D., Mather, A.E., and Otto, F.D, "Solubility of acid gases in methyldiethanolamine solutions," presented to the CGPA Association Meeting, Calgary, Alberta, Mar. 12, 1981.
  8. Jou, F.Y., Carroll, J.J., Mather, A.E., and Otto, F.D., "Solubility of carbon dioxide and hydrogen sulfide in a 35 wt % aqueous solution of methyldieth anolamine," Canadian Journal of Chemical Engineering, Vol. 71, 1993.
  9. Lidal, H., "Carbon dioxide removal in gas treating processes," PhD Thesis, the University of Trondheim, Norwegian Institute of Technology, Department of Chemical Engineering, 1992.
  10. Knudsen, B., Tjelle, S., and Linga, H., "A new approach towards environmentally friendly desulphurisation," presented to the SPE Sixth International Conference on Health, Safety & Environ ment in Oil and Gas Exploration, Kuala Lumpur, Mar. 20-22, 2002.

The authors

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Finn P. Nilsen is a mechanical engineer and managing director for Framo Purification AS, Bergen, Norway. For the past 4 years, he has worked in developing new technologies for gas purification. Previously he worked for Christian Michelsen Research, Bergen, and Kongsberg Defence & Aerospace, Kongsberg, Norway. Nilsen is a member of SPE and GPA.

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Inge Sverre Lund Nilsen is an application engineer with Framo Purification AS, Bergen, Norway. His background includes natural gas purification, gas processing, and combustion engineering. He received an MSc in mechanical engineering from the Norwegian Institute of Technology, Trondheim, Norway.

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Haavard Lidal is a specialist with Statoil AS, Oslo. He is a technical manager of the field installations of the Kristin project. He has 10 years' experience with Statoil on several of the largest North Sea and Norwegian Sea projects. Lidal received an MSc and PhD in chemical engineering from the Norwegian Institute of Technology, Trondheim, Norway. He is a member of AIChE and GPA.