PROCESSING
New NGL-recovery process
provides viable alternative
Robert R. Huebel
Michael G. Malsam
Randall Gas Technologies
Houston
Operational scenarios for two uses of a
new refrigeration process for recovering NGLs from natural gas have shown
it to enhance operability and reduce
capital and operating expenditures
when compared with the two more
traditional process choices—straight
refrigeration and turboexpander.
Straight refrigeration units that
most often use propane as refrigerant
have proven to be economical and reliable. Their operating temperature,
however, typically about –35° F., limits
NGL extraction. For higher NGL recovery, today’s processor is left with a
cryogenic turboexpander.
IPOR (IsoPressure Open Refrigeration) has been developed by Randall
Gas Technologies, a division of Lummus Technology, a CB&I company,
to bridge this gap. The advanced refrigeration process can economically
achieve essentially total C3+ recovery
from most natural gas streams. Using
conventional closed-loop mechanical refrigeration combined with an
open-loop mixed refrigeration cycle, the
CASE STUDY 1:
PLANT DESIGN BASIS
Table 1
Feed gas:
Flow, MMscfd
100
Pressure, psig 70
Temperature, oF. 80
Composition, mol %:
N2 & CO2
1.8
C1
75.0
C2
16.1
C3+
7.3
Residue gas:
Pressure, psig 1,200
Heating value, btu/scf max 1,100
NGL product specifications:
C2/C3 liquid volume ratio
0.02
CASE STUDY 2:
PLANT DESIGN BASIS
Table 2
Based on a presentation to the GPA Europe Annual Conference,
Prague, Sept. 21-23, 2011.
Diverse environments
Natural gas conditioning and processing plants are somewhat unique in that
the raw material feedstock is typically
fed into the plant at the pressure, flow
rate, and composition at which it is
produced.
Consequently, natural gas processing plants have considerable variation
in size, complexity, and configuration,
depending upon specific reservoir
production characteristics, geography,
customer specifications, and market
drivers. These range from simple dewpoint plants with capacities less than 5 MMscfd and minimal hydrocarbon recovery to large deep cut ethane extraction straddle plants
which process in excess of 1 bcfd .
With such a diverse operating environment, it is a bit sur-
Feed gas:
Flow, MMscfd
20
Pressure, psig 200
Temperature, oF. 50
Composition, mol %:
N2 & CO2
1.9
C1
81.2
C2
9.3
C3+
7.6
Residue gas:
Pressure, psig 950
Hydrocarbon dewpoint, oF.
–5
NGL product specifications:
C2/C3 liquid volume ratio
0.02
new technology
can achieve NGL recovery efficiencies
comparable to that of advanced turboexpander cycles but for lower capital
and operating expenditures.
This article reviews the fundamentals of the IPOR process, including
process features, benefits, and applicability. It also presents case studies
that compare process performance
with both straight refrigeration and
advanced turboexpander cycles and
economic analysis.
PROCESSING
IPOR PROCESS FOR HIGH LPG RECOVERY
Feed
FIG. 1
Propane refrigeration
compressor
Sales
Gas-gas
exchanger
Mixed refrigerant
compressor
Mixed refrigerant
gas-gas exchanger
De-ethanizer
overhead
condenser
De-ethanizer
reflux drum
De-ethanizer
overhead
separator
De-ethanizer
reboiler
NGL
prising that natural gas processors have had essentially only
two process technology choices for extracting hydrocarbon
liquids from natural gas: either straight refrigeration or turboexpander. Among more than 1,600 operating natural gas
processing plants shown in Oil & Gas Journal’s Worldwide
Gas Plant Survey, about 80% use either straight refrigeration
or turboexpander technology (OGJ, June 6, 2011, p. 88).
With the last new lean oil plant built some 30 years ago,
the estimated portion of new gas plants built today using
these two technologies is greater than 95%.
Straight refrigeration units, which most often use propane or ammonia, can be built for essentially any capacity
or feed-gas composition, are of mild steel construction, are
relatively simple to construct and operate, and have proven
to be economical and reliable. However, with their operating
temperature typically limited to about –35° F., their capability for NGL extraction is limited.
For higher NGL recovery, today’s processor has but a
single choice: cryogenic turboexpander. Since its inception
in the late 1960s, turboexpander technology has evolved
into the technology of choice for deep NGL-product recovery. As designs were refined, turboexpander technology
essentially displaced lean-oil technology for high LPG or
ethane-extraction applications.
Several variations of the technology are available, depending upon the targeted product recovery and feed-gas conditions, with proprietary designs offering even higher efficiencies. With operating temperatures as low as –200° F., NGL
product recoveries approaching 98%+ are technically feasible.
With straight refrigeration technology, the benefits for
the customer include low capital and operating expenditures
(CAPEX and OPEX), a broad range of applicability, early
production capabilities, but limited NGL recovery. Expander
technology offers superior NGL-recovery potential but high-
z1120109OGJphu01
De-ethanizer
PROCESSING
TURBOEXPANDER PROCESS
FIG. 2
Residue gas to
recompression
(160 psig)
–129° F.
Fuel
Cold
separator
Demethanizer
Booster
compressorexpander
De-ethanizer
De-ethanizer
reboiler
Feed
(400 psig)
–58° F.
NGL
Reboilers
Propane
refrigerant
–35° F.
er CAPEX and OPEX and a longer time to initial operation
due to the long lead time of such specialty equipment as the
turboexpander and brazed aluminum heat exchangers.
Ethane-rich cycle
The advanced refrigeration NGL extraction process can economically achieve deep NGL extraction from most natural
gas streams. Using conventional closed-loop mechanical refrigeration combined with an open-loop mixed refrigeration
cycle, this process can provide performance comparable to
that of advanced turboexpander technologies but with much
lower CAPEX and OPEX.
Unique about the IPOR process is its open-loop ethanerich mixed refrigeration cycle. This refrigerant, extracted
from the feed gas itself, is a mixture of predominantly ethane
with lower concentrations of methane, propane, and other
feed-gas constituents.
This refrigeration cycle serves a dual purpose: producing the
cryogenic refrigeration for the process to enable lower temperature operation while at the same time providing a reflux stream
to the fractionation column, the combination of which produces high product extraction and thermal efficiencies.
The extraction process can be configured in several ways,
depending on the feed stream, site conditions, and project
objectives. Fig. 1 depicts one configuration of the IPOR technology recommended for high recovery LPG applications.
Feed gas, at a pressure typically 300-550 psig, is initially
cooled and partially condensed in the gas-gas heat exchanger by cross exchange with cold residue gas and propane refrigerant. A conventional brazed aluminum heat exchanger
appears in the flow diagram; however, shell-and-tube exchangers can also be used for this service.
The cooled and partially condensed feed-gas stream is
then fed to the middle section of the de-ethanizer, which
uses either trays or packing or a combination of these to
effect the desired product separation. Below the feed tray,
the stripping section of the column selectively removes the
lighter fractions to meet product specifications, which normally is 2-5% ethane in the recovered propane. Heat for the
separation is provided by the de-ethanizer reboiler, which is
a conventional shell-and-tube heat exchanger, with the heat
supplied from the plant heating medium system.
In the upper section of the de-ethanizer, above the feed
tray, the cooled feed gas flows counter-currently to the reflux
stream, which is fed to the top tray in a conventional manner. The reflux provides additional cooling for the feed-gas
stream and also selectively absorbs the propane and heavier
components from the gas, thereby providing high product
recovery efficiencies.
The overhead gas stream from the de-ethanizer, at this
point in the process containing primarily the light ends from
the feed-gas stream and a small portion of the propane, is
z120109OGJphu02
Residue
recycle
PROCESSING
IPOR PROCESS
Feed
FIG. 3
Propane refrigerant
compressor
Sales
120 psig
Mixed refrigerant
compressor
365 psig
410 psig
–10˚ F.
Gas-gas
exchanger
Mixed refrigerant
gas-gas exchanger
Fuel
De-ethanizer
overhead
condenser
–42˚ F.
De-ethanizer
reflux drum
De-ethanizer
De-ethanizer
overhead
separator
De-ethanizer
reboiler
NGL
further cooled in the de-ethanizer overhead condenser by
cross exchange with cold residue gas and the ethane-rich
mixed refrigerant stream.
The cooled and partially condensed gas stream flows to
the de-ethanizer overhead separator. The liquid from this
separation, a mixture of methane, ethane, and propane, is
used as the refrigerant for the open-loop mixed refrigerant
cycle. The de-ethanizer overhead separator therefore has a
twofold function: It acts as a conventional two-phase gasliquid separator, and it provides surge capacity for the liquid
mixed refrigerant system.
From the de-ethanizer overhead separator, the pressure
of the liquid mixed refrigerant is reduced, creating a JouleThomson refrigeration effect: This cold stream provides the
desired cooling in the de-ethanizer overhead condenser. The
pressure of the low-pressure mixed refrigerant, usually in
the range of 100-200 psig, is selected to satisfy the cooling
requirements in the de-ethanizer overhead condenser and to
minimize the compression power requirements.
From the de-ethanizer overhead condenser, the mixed
refrigerant stream is heated further as it flows through the
mixed refrigerant gas-gas exchanger to the mixed refrigerant compressor. The discharge pressure of this compressor
is normally about 40 psig higher than the operating pressure
of the de-ethanizer.
The mixed refrigerant compressor is of conventional design
and can be either reciprocating, centrifugal, or screw type, depending upon project requirements and customer preferences. Drivers may be gas turbine, gas engine, or electric motor.
The compressor can be packaged with driver, scrubbers, and
discharge cooler following standard industry practice.
The compressed ethane-rich, mixed refrigerant stream is
then cooled and partially condensed in the mixed refrigerant gas-gas exchanger. Cooling for this exchanger is provid-
z1120109OGJphu03
–75˚ F.
PROCESSING
MECHANICAL REFRIGERATION
FIG. 4
Residue gas
Gas-refrigerant
exchanger
Cooler
Recompressor
225 psig
Fee gas
(975 psig)
Gas-gas
exchanger
Gas chiller
Cold
separator
Reboiler
C3+ product
ed by low-temperature mixed refrigerant and propane. The
two-phase stream flows to the de-ethanizer reflux drum, a
conventional two-phase gas liquid separator. This liquid is
used to provide reflux to the de-ethanizer column, thereby
completing the “open” cycle of the mixed refrigerant loop.
Noncondensable vapors, consisting mainly of methane,
are directed back into the process via the de-ethanizer overhead separator and eventually exit the process into the residue gas stream or may be used as fuel.
The closed-loop propane refrigeration is of conventional
natural gas industry design and construction. In a typical
CASE STUDY 1: RESULTS
Table 3
Feed-gas capacity: 100 MMscfd
IPOR
process
Product recovery:
C3
C4+
NGL production, b/d
99.5
100.0
98.8
100.0
5,173
5,131
Power, bhp:
Inlet compression
9,780
Residue compression
7,060
Refrigeration
4,730
Pumps, air coolers
490
Total power
22,020 10,460
11,930
1,910
560
24,860
Gas compression
Process compression
Turboexpander
–16,000
6,020
–16,000
8,860
Major equipment count—process:
Turboexpander
––
Pumps
––
Columns
1
All other
24
Total major equipment count—process
25
1
4
2
24
31
IPOR process, the process refrigeration temperature is in the
range of –10° F. to –20° F.; other refrigerants, therefore, such
as ammonia may be used as well.
For the LPG-recovery configuration above, product extraction efficiencies are excellent, with C3 recovery in the range of
95-99%+, with essentially 100% recovery of the C4+ fraction.
From a thermal efficiency perspective, the IPOR process
requires about 15-40% less compression power than a comparable turboexpander design. As a result, plants using the
IPOR technology will also have lower emissions and a smaller carbon footprint.
CASE STUDY 1: ECONOMIC ANALYSIS
Feed-gas capacity: 100 MMscfd
IPOR
process, $
Table 4
Turboexpander
CAPEX
<11.0 million
Base
Savings
Turboexpander
Stainless steel pumps (2)
Carbon steel pumps (2)
Stainless steel column
Compression
OPEX
<700,000/year
Base
Fuel savings @ $4.50/MMbtu
NGL revenue
>220,000/year
Base
Shrinkage @ $4.50/MMbtu
Crude @ $80/bbl
C3 @ 60% of crude
Trans. & frac. @ $0.05/gal
96% availability
z120109OGJphu04
Stabilizer
–10° F.
PROCESSING
IPOR PROCESS FOR CASE STUDY 2
Feed
FIG. 5
Residue gas
to recompressor
Propane refrigerant
compressor
125 psig
Mixed refrigerant
compressor
435 psig
410 psig
Gas-gas
exchanger
Mixed refrigerant
gas-gas exchanger
–20° F.
Gas chiller
Gas chiller
De-ethanizer
overhead
condenser
De-ethanizer
reflux drum
De-ethanizer
De-ethanizer
reboiler
De-ethanizer
overhead
separator
NGL
The process utilizes equipment and materials that are all
well proven within the natural gas processing industry. Most
of the unit can be of carbon steel or low-temperature carbon
steel construction; typically the only major equipment item
that requires stainless steel construction is the de-ethanizer
overhead separator.
The only rotating equipment required for the IPOR process
is the refrigerant compressor. The process requires no cryogenic turboexpander or light hydrocarbon pumps. As a result:
• Reliability and operability will be comparable to that
of a conventional refrigeration process and should exceed
that of a modern day turboexpander facility, given the fewer
items of rotating equipment.
• The process offers superior economics for almost any
feed-gas rate, from as low as 5 MMscfd to 1 bcfd+.
• Almost infinite turndown capacity is possible with an
IPOR process, to feed-gas rates as low as 10% of design, lim-
ited only by the performance of in-line control instruments,
i.e., control valves, meters, etc., unlike turboexpander designs,
which suffer from an inherent loss of efficiency at reduced flows.
The process can be designed for a wide variety of feed-gas
compositions, site conditions, and capacities. Ethane recovery can be incorporated into an IPOR process design, with
ethane recoveries up to 80%, depending upon feed-gas composition. Equipment can be incorporated to allow for future
ethane recovery, or the initial design can permit operation in
ethane-rejection/ethane-recovery mode.
The process was developed based on proven technologies and equipment employed extensively in gas plants. All
the equipment incorporated into the process design is well
within the natural gas industry’s experience and capability.
The low equipment count, small footprint, and process simplicity of the technology permit a compact layout and a high
degree of modularization.
z1120109OGJphu05
–105° F.
PROCESSING
CASE STUDY 2: RESULTS
Feed-gas capacity: 20 MMscfd
Table 5
IPOR
process
CASE STUDY 2: ECONOMIC ANALYSIS
Feed-gas capacity: 20 MMscfd
Refrigeration
Product recovery, %:
C3
99.0
C4
100.0
C5+
100.0
NGL production, b/d
1,087
Power, bhp
Inlet compression
810
Residue compression
975
Refrigeration
915
Total
2,700
Facility major equipment count
44
Refrigeration
CAPEX
>2.0 million
Mole-sieve dehy vs. glycol
Additional plate fins
Additional compression
OPEX (fuel)
>150,000/year
Fuel value @ $4.50/MMbtu
NGL revenue
>5.1 million/year
Shrinkage @ $4.50/MMbtu
Crude @ $80/bbl
C3 @ 60% of crude
C4 @ 80% of crude
C5 @ 90% of crude
Trans. & frac. @ $0.05/gal
96% availability
Internal rate of return
155%
38% tax rate
Double-declining balance
depreciation rate
20-year plant life
Payback period
<6 months
33.1
60.0
83.4
532
1,830
270
2,100
45
Marcellus plant
A recent study compared the IPOR process with modern turboexpander technology. Feedstock for the new plant is from
the Marcellus shale, a region with limited existing oil and
gas infrastructure and no existing ethane market. Demand
for LPG in the region is strong, with extracted LPG sold into
the local market.
As a result, the customer wanted to maximize LPG production. Due to the richness of the gas, some ethane extraction was required to meet the residue-gas pipeline specifications, with the ethane consumed within the plant as fuel.
The field’s gathering system operated at low pressure, with
residue gas delivered into an existing high pressure pipeline.
Table 1 summarizes the design basis for the plant.
Two process technologies were evaluated: conventional
turboexpander and the IPOR process.
The turboexpander process utilized in the study was a
modern design (Fig. 2). Due to the richness of the feed gas,
a propane refrigeration system with a low stage operating
temperature of –35° F. at 3.4 psig was integrated into the
process design to provide supplemental cooling. A portion
of the ethane vapor stream from the overhead of the de-eth-
Table 6
IPOR
process, $
Base
Base
Base
Base
anizer column is consumed as fuel to achieve the residue-gas
heating value specifications.
Fig. 3 illustrates the IPOR process used in the study. Feed
gas enters the process unit at a compressor interstage pressure of about 365 psig. The propane refrigeration system operates at –10° F., much warmer than that required by the turboexpander process, and 16.7 psig. The minimum operating
temperature of the de-ethanizer column is –42° F. and is of
low-temperature carbon steel construction.
To achieve the residue-gas pipeline heating value specification, a portion of the ethane-rich noncondensable vapors
from the de-ethanizer reflux drum is consumed as fuel, with
the remainder mixing with the residue gas via the de-ethanizer overhead separator.
Tables 3 and 4 summarize the results of the study.
Compared with the turboexpander design, the IPOR
PROCESS COMPARISONS
Base
Table 7
Refrigeration
IPOR process
Turboexpander
Applicability
Feed-gas volume, MMscfd
Feed-gas pressure, psig
Feed-gas hydrocarbons
Any
Any
Lean-Rich
Any
<600
Moderate-Rich
50+
Any
Lean-Rich
NGL recovery
Ethane, %
Propane, %
Butane, %
Gasoline, %
N/A
20-40
50-70
70-90
40-80
99.9
100
100
95
99.9
100
100
Constructability
Materials of construction
Modularization potential
Long lead equipment delivery, months
Carbon steel
High
4-6
Limited alloy
High
4-6
Extensive alloy
High
8-12
Operability
Turndown, % of design
Reliability
Maintenance
10
High
Low
10
High
Low
50
High
Medium
Economics
CAPEX
OPEX
Low
Low
Medium
Medium
High
High
PROCESSING
process:
the feed-gas compression, with residue gas sent directly to
1. Achieves higher NGL recovery.
the sales gas pipeline.
2. Requires about 32% less process compression power.
Fig. 5 illustrates the IPOR process used in the study. Feed
3. Requires about 20% less major equipment.
gas enters the IPOR unit at a compressor interstage pressure
4. Requires less rotating equipment.
of about 410 psig. For this design, the gas-gas exchangers
As a result, economics of the IPOR process are clearly suand chillers were conventional shell-and-tube design. All
perior to the turboexpander design, both from an OPEX and
of the noncondensable vapors from the de-ethanizer reflux
a CAPEX perspective (Table 4). Estimated capital cost of the
drum flow to the de-ethanizer overhead separator and on to
IPOR process design was $11 million less than that of the
the residue gas stream.
turboexpander plant, the savings being the result of:
Tables 5 and 6 summarize the study’s results.
1. Less installed compression.
From these results, key observations include the following:
2. No turboexpander.
1. NGL production with the IPOR unit is more than dou3. No light hydrocarbon/cryogenic pumps.
ble that of the refrigeration plant.
4. No stainless steel demethanizer column.
2. Complexity of the two designs is comparable,
5. Less alloy material.
From an operating cost perspective, the IPOR process was estimated to consume about $700,000/
year less in utilities, the savings
elson-farrar cost indexes
resulting from lower compression
power requirements, and hence
Refinery construction (1946 basis)
(Explained in OGJ, Dec. 30, 1985, p. 145, and at www.pennenergy.com/index/research-and_data/oil-and_gas/Statisticfuel gas consumption.
Definitions.html; click “Nelson-Farrar Cost Indices”)
N
Northwest Canada
A second study was recently completed comparing the IPOR process
to a straight refrigeration process.
Location for this plant is in northwest Canada, an area of existing oil
and gas production but no NGL or
ethane pipeline infrastructure. Liquids produced in the plant would be
trucked to market.
The primary objective of the customer in this application was to deliver a marketable sales gas. Given
the current favorable economic climate for gas liquids, however, incremental LPG recovery was of interest
if economical.
The basis of design of the plant
for the study is discussed below. Table 2 summarizes the design basis.
The straight refrigeration process used in the study was a traditional design (Fig. 4), with process
temperature selected to achieve the
pipeline dewpoint specification,
thereby minimizing both CAPEX
and OPEX. Propane was used as
the refrigerant, with glycol injection
used for hydrate inhibition and dehydration. Feed gas for the refrigeration unit was taken downstream of
1962
Pumps, compressors, etc.
222.5
Electrical machinery
189.5
Internal-comb. engines
183.4
Instruments
214.8
Heat exchangers
183.6
Misc. equip. average
198.8
Materials component
205.9
Labor component
258.8
Refinery (Inflation) Index
237.6
1980
2008
2009
2010
Sept.
2010
Aug.
2011
Sept.
2011
777.3
1,949.8
2,011.4
2,030.7
2,036.4
2,119.6
2,120.5
394.7
515.6
515.5
513.9
513.7
515.0
514.1
512.6
990.9
1,023.0
1,027.8
1,021.2
1,036.3
1,036.3
587.3
1,342.1
1,394.8
1,435.1
1,437.3
1,458.2
1,461.6
618.7
1,354.6
1,253.8
1,116.0
1,103.5
1,103.5
1,253.8
578.1
1,230.6
1,239.7
1,224.7
1,222.4
1,246.5
1,277.3
629.2
1,572.0
1,324.8
1,480.1
1,489.4
1,619.7
1,627.6
951.9
2,704.3
2,813.0
2,909.3
2,923.3
2,992.9
3,000.2
822.8
2,251.4
2,217.7
2,337.6
2,349.8
2,443.6
2,451.2
Refinery operating (1956 basis)
(Explained in OGJ, Dec. 30, 1985, p. 145, and at www.pennenergy.com/index/research-and_data/oil-and_gas/StatisticDefinitions.html; click “Nelson-Farrar Cost Indices”)
1962
2009
2010
Sept.
2010
Aug.
2011
Sept.
2011
1,951.3
978.5
1,184.9
1,048.8
1,267.6
1,196.5
237.9
264.5
281.7
277.9
255.1
262.4
439.9
1,092.2
1,177.1
1,279.4
1,289.0
1,270.6
1,267.5
226.3
460.8
445.2
454.5
463.9
498.2
483.0
1980
2008
100.9
810.5
93.9
200.5
123.9
131.8
Invest., maint., etc.
121.7
Chemical costs
96.7
Fuel cost
Labor cost
Wages
Productivity
Operating indexes
Refinery
103.7
Process units*
103.6
324.8
830.8
812.4
850
854.5
888.6
891.3
229.2
472.5
406.2
449.8
444.0
557.9
560.1
312.7
674.1
582.6
628.2
615.9
652.2
650.0
457.5
1,045.1
706.1
796.8
749.5
831.2
809.5
*Add separate index(es) for chemicals, if any are used. See current Quarterly Costimating in first issues for January,
April, July, and October. These indexes are published in the first of each month. They are compiled by Gary Farrar, OGJ
Contributing Editor. Indexes of selected individual items of equipment and materials are also published on the Costimating
page in first issues for January, April, July, and October.
PROCESSING
based upon major equipment count, which should result in similar operability and reliability. (Major equipment count in this case includes the entire plant facility, including dehydration, utilities, and off sites.)
As a result, economics of the IPOR process are once again
superior to the refrigeration unit, taking into account incremental differences in both OPEX and CAPEX (Table 6). Estimated capital cost of the IPOR process design was $2 million
more than that of the refrigeration plant, the additional cost
being the result of:
1. More installed compression.
2. Additional heat-exchanger costs.
3. Additional cost of the molecular-sieve dehydration
system vs. the glycol injection system utilized in the refrigeration plant design.
4. More alloy materials.
The operating cost of the IPOR process was estimated to
be about $150,000/year more than the refrigeration process.
The additional cost was primarily the result of the higher
compression power requirements of the IPOR process, and
therefore more fuel-gas consumption.
NGL production with the IPOR unit is more than double
that of the refrigeration plant. The value of this additional NGL revenue was estimated at $5.1 million/year. Based
upon the economic assumptions itemized in Table 6, the calculated internal rate of return of the IPOR plant investment
is 155%, with a payback of fewer than 6 months.
While the IPOR unit requires somewhat more CAPEX
and OPEX than a “minimal type investment” of the refrigeration unit, these costs are more than compensated for with
the increased NGL revenue.
Table 7 summarizes the comparisons discussed in this
article.
References
1. Malsam, Michael G, “IPOR Technology—A new means
of LPG recovery,” Gas Processors Association Annual Convention, “High Definition at 90—Advancing the Midstream
Vision,” March 2011.
2. “Gas Processing with Cryogenic Turboexpander Technology,” Randall Gas Technologies, Houston; January 2011 Edition.
The authors
Robert R. Huebel (rhuebel@cbi.com) is
vice-president of technology of Randall Gas
Technologies, a division of Lummus Technology Inc., a CB&I company. His previous role
was as president of the ABB Randall Corp.
He joined the company in 1976. Huebel
has more than 40 years’ experience in the
domestic and international engineering, procurement, and construction and natural gas
processing industries, including process engineering, project
management, contract development, and executive management. He holds a BSc in chemical engineering and an MBA
from the University of Houston. Huebel is a registered professional engineer in seven states, is a member of the American
Institute of Chemical Engineers and Project Management
Institute, and currently serves on the board of directors of the
Gas Processors Suppliers Association.
Michael G. Malsam (mmalsam@cbi.com) is
senior principal process engineering specialist
for Randall Gas Technologies, which he joined
in 1998. He has more than 30 years’ experience in the domestic and international EPC
and natural gas processing, including process
engineering, project development, and project
management. Malsam holds a BSc in chemical and petroleum refining engineering from the
Colorado School of Mines. He is a member of the American Institute of Chemical Engineers and the American Chemical Society.
Reprinted with revisions to format, from the January 9, 2012 edition of Oil & Gas Journal
Copyright 2012 by PennWell Corporation