National Renewable Energy Laboratory 

Innovation for Our Energy Future 

A national laboratory of the U.S. Department of Energy

Office of Energy Efficiency & Renewable Energy

NREL is operated by Midwest Research Institute ● Battelle Contract No. DE-AC36-99-GO10337 

Equipment Design and Cost 

Estimation for Small Modular 

Biomass Systems, Synthesis 

Gas Cleanup, and Oxygen 

Separation Equipment 

Task 1: Cost Estimates of Small 

Modular Systems 

Nexant Inc. 

San Francisco, California 

Subcontract Report 

NREL/SR-510-39943 

May 2006 
Equipment Design and Cost 

Estimation for Small Modular 

Biomass Systems, Synthesis 

Gas Cleanup, and Oxygen 

Separation Equipment 

Task 1: Cost Estimates of Small 

Modular Systems 

Nexant Inc. 

San Francisco, California 

NREL Technical Monitor: Kelly Ibsen 

Prepared under Subcontract No. ACO-5-44027 

Subcontract Report 

NREL/SR-510-39943 

May 2006 

National Renewable Energy Laboratory 

1617 Cole Boulevard, Golden, Colorado 80401-3393 

303-275-3000 • www.nrel.gov 

Operated for the U.S. Department of Energy 

Office of Energy Efficiency and Renewable Energy 

by Midwest Research Institute • Battelle 

Contract No. DE-AC36-99-GO10337
This publication was reproduced from the best available copy 

Submitted by the subcontractor and received no editorial review at NREL

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Contents 

Section Page 

Executive Summary ..............................................................................................................................ES-1

Introduction and Methodology..................................................................................................................1

Section 1 Fluidized Catalytic Cracking ............................................................................................1-1

1.1 Introduction and Methodology .................................................................................................1-1 

1.2 Technology Background...........................................................................................................1-2 

1.3 Process Descriptions and Design Assumptions........................................................................1-3 

1.3.1 Large Scale (Side-by-Side) FCC .....................................................................................1-8 

1.3.2 Small Scale (Stacked) FCC............................................................................................1-10 

1.4 Operating Costs and Utility Requirements.............................................................................1-12 

Section 2 Natural Gas Reforming .....................................................................................................2-1

2.1 Introduction and Methodology .................................................................................................2-1 

2.2 Technology Background...........................................................................................................2-1 

2.3 Process Descriptions and Design Assumptions........................................................................2-1 

2.3.1 Small Scale SMR Design.................................................................................................2-2 

2.3.2 Large Scale SMR Design.................................................................................................2-8 

2.4 Operating Costs and Utility Requirements.............................................................................2-14 

Section 3 Natural Gas Liquid Expanders.........................................................................................3-1

3.1 Introduction and Methodology .................................................................................................3-1 

3.2 Technology Background...........................................................................................................3-1 

3.3 Process Descriptions and Design Assumptions........................................................................3-3 

3.3.1 Small Scale NGL Expander Design.................................................................................3-3 

3.3.2 Large Scale NGL Expander Design.................................................................................3-9 

3.4 Operating Costs and Utility Requirements.............................................................................3-14 

Section 4 Summary and Conclusions................................................................................................4-1

Appendix A Technology Evaluations....................................................................................................A-1

A.1 Refinery Processes...................................................................................................................A-1 

A.2 Claus Sulfur Recovery.............................................................................................................A-5 

A.3 Natural Gas Reforming............................................................................................................A-7 

A.4 Methanol from Natural Gas.....................................................................................................A-9 

A.5 Ammonia Production.............................................................................................................A-11 

A.6 Gas-To-Liquids......................................................................................................................A-13 

A.7 Natural Gas Liquids Expanders.............................................................................................A-13 

A.8 Cryogenic Oxygen Plants......................................................................................................A-16 

A.9 Ethylene and other Olefins Production..................................................................................A-18

Task 1 Final Report ii 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Figures and Tables 

Figure Page 

1-1 Large Scale FCC Design...............................................................................................................1-4 

1-2 Small Scale FCC Design...............................................................................................................1-5 

2-1 Small Scale SMR Process Flow Diagram.....................................................................................2-3 

2-2 Typical Small Scale SMR Reactor Design ...................................................................................2-7 

2-3 Large Scale SMR Process Flow Diagram...................................................................................2-10

2-4 Typical Large Scale SMR Reactor Design .................................................................................2-13

2-5 Large Scale SMR Reactor Heat Integration................................................................................2-13 

3-1 Small-Scale NGL Expander Process Flow Diagram ....................................................................3-4 

3-2 Large-Scale NGL Expander Process Flow Diagram ..................................................................3-11 

4-1 SMR Project Costs, Broad Size Range .........................................................................................4-1 

A-1 2004 US Refinery Sizes...............................................................................................................A-2 

A-2 Typical Claus Plant BFD ............................................................................................................A-6 

A-3 Steam Methane Reformer Process Flow, Ammonia Plant...........................................................A-8 

A-4 Low Pressure Methanol Process Flow Diagram........................................................................A-10 

A-5 Ammonia Production Capacity and Energy Consumption (1960-2010) ...................................A-11 

A-6 Ammonia Production Flowsheet................................................................................................A-12 

A-7 Schematic for Neptune II Turboexpander Plant ........................................................................A-15 

A-8 Generic Cryogenic Oxygen Plant Flow Diagram ......................................................................A-17 

A-9 KBR ScoreTM (Selective Cracking Optimum Recovery) Ethylene Technology........................A-19 

A-10 Ethylene Plant Size Progression, 1945-2005 .............................................................................A-20 

Task 1 Final Report iii 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Table Page 

A Total FCC Project Costs ............................................................................................................ ES-1 

B Total Steam Methane Reformer Project Costs........................................................................... ES-1 

C Total Natural Gas Expander Project Costs ................................................................................ ES-2 

D Design Ratios and Predicted Costs, All Technologies............................................................... ES-2 

1-1 Large Scale FCC Stream Properties..............................................................................................1-4 

1-2 Small Scale FCC Stream Properties..............................................................................................1-5 

1-3 Large Scale FCC Equipment List .................................................................................................1-6 

1-4 Small Scale FCC Equipment List .................................................................................................1-7 

1-5 Total FCC Project Costs, 2002 Basis............................................................................................1-8 

1-6 FCC Operating Costs and Utility Requirements.........................................................................1-12 

2-1 SMR Project Costs, 2005 Basis ....................................................................................................2-2 

2-2 Small Scale SMR Stream Properties.............................................................................................2-4 

2-3 Small Scale SMR Equipment List ................................................................................................2-6 

2-4 Large Scale SMR Stream Properties...........................................................................................2-11 

2-5 Large Scale SMR Equipment List ..............................................................................................2-12 

2-6 Major Process Conditions, Large and Small SMR Designs .......................................................2-14 

2-7 SMR Operating Costs And Utility Requirements.......................................................................2-14 

3-1 Typical NGL Feedstock Compositions.........................................................................................3-2 

3-2 NGL Expander Project Costs, 2005 Basis ....................................................................................3-3 

3-3 Small Scale Expander Stream Properties......................................................................................3-5 

3-4 Small-Scale NGL Expander Equipment List................................................................................3-7 

3-5 Large Scale Expander Stream Properties....................................................................................3-12 

3-6 Large-Scale NGL Expander Equipment List..............................................................................3-13 

3-7 NGL Expander Operating Costs And Utility Requirements.......................................................3-14 

4-1 Design Ratios and Predicted Costs, All Technologies..................................................................4-2 

A-1 Largest/Smallest US Refineries ...................................................................................................A-2 

A-2 Atmospheric/Vacuum Crude Distillation Unit Capacities...........................................................A-3 

A-3 Alkylation Unit Capacities...........................................................................................................A-3 

A-4 Aromatics Unit Capacities ...........................................................................................................A-4 

A-5 Fluidized Catalytic Cracking Unit Capacities (Fresh Feed) ........................................................A-4 

A-6 Hydrocracking Unit Capacities....................................................................................................A-4 

A-7 Naphtha Reforming Unit Capacities............................................................................................A-5 

A-8 Desulfurization Unit Capacities...................................................................................................A-5 

A-9 Claus Sulfur Recovery Capacities ...............................................................................................A-7 

Task 1 Final Report iv 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Table Page 

A-10 SMR Plant Capacities ..................................................................................................................A-9 

A-11 Methanol Plant Capacities .........................................................................................................A-10 

A-12 Ammonia Plant Capacities.........................................................................................................A-13 

A-13 Typical NGL Feedstock Compositions......................................................................................A-14 

A-14 NGL Turboexpander Plant Capacities .......................................................................................A-16 

A-15 Cryogenic Oxygen Plant Capacities ..........................................................................................A-17 

A-16 Ethylene Plant Capacities ..........................................................................................................A-20 

Task 1 Final Report v 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Executive Summary 

This deliverable is the Final Report for Task 1, Cost Estimates of Small Modular Systems, as part of 

NREL Award ACO-5-44027, “Equipment Design and Cost Estimation for Small Modular Biomass 

Systems, Synthesis Gas Cleanup and Oxygen Separation Equipment”. Subtask 1.1 looked into processes 

and technologies that have been commercially built at both large and small scales, with three 

technologies, Fluidized Catalytic Cracking (FCC) of refinery gas oil, Steam Methane Reforming (SMR) 

of Natural Gas, and Natural Gas Liquids (NGL) Expanders, chosen for further investigation. These 

technologies were chosen due to their applicability relative to other technologies being considered by 

NREL for future commercial applications, such as indirect gasification and fluidized bed tar cracking. 

Research in this subject is driven by an interest in the impact that scaling has on the cost and major 

process unit designs for commercial technologies. Conclusions from the evaluations performed could be 

applied to other technologies being considered for modular or skid-mounted applications. 

Subtask 1.2.1 presented PFDs and preliminary material/energy balances for each of the designs. Subtasks 

1.2.2 through 1.2.4 built upon the information presented in Subtask 1.2.1 by providing equipment 

specification sheets, capital and installation costs, cost breakdowns for individual pieces of equipment, 

and operating costs. The following deliverable provides detail for large and small scale designs of each 

technology. 

For each technology, a basis was established for throughput and cost at each scale. Results from the 

technical analysis can be seen in Tables A through C: 

TABLE A TOTAL FCC PROJECT COSTS 

Large Scale Plant Small Scale Plant 

Gas Oil Feed Rate (BPSD) 62,000 2,500 

Reactor/Regen TPI ($MM) 102.6 11.0 

Total Project Investment ($MM) 191.6 20.5 

Total Investment/BBL Feed ($) ~3,000 ~8,200 

TABLE B TOTAL STEAM METHANE REFORMER PROJECT COSTS 

Large Scale Plant Small Scale Plant 

Natural Gas Feed Rate (MMSCFD) 167.8 0.024 

Installed Cost ($MM) 244.0 0.26 

Total Plant Investment ($MM) 368.5 0.39 

Total Investment/SCFD Feed ($) 2.2 16.2 

Task 1 Final Report ES-1 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
TABLE C TOTAL NATURAL GAS EXPANDER PROJECT COSTS 

Large Scale Plant Small Scale Plant 

Natural Gas Feed Rate (MMSCFD) 400 75 

Installed Cost ($MM) 46.4 12.0 

Total Plant Investment ($MM) 62.4 16.1 

Total Investment/SCFD Feed ($) 0.16 0.21 

Although the technologies considered and the relative ranges between the large and small scale designs 

were all different, a common relationship existed between the large and small scale designs that could be 

applied to other technologies. Using a simple power relationship defined by the following, 

7.0

Capacity Scale Large

Scale Small Capacity Cost Scale Small CostScale Large ⎟

⎟

⎠

⎞

⎜

⎜

⎝

⎛ = ∗

it was determined that the relationship between each case could be predicted with some degree of 

accuracy. Table D below shows the size and cost ratios for each technology, along with the predictive 

capacity of using a power relationship. 

TABLE D DESIGN RATIOS AND PREDICTED COSTS, ALL TECHNOLOGIES 

FCC SMR NG Expander 

Large/Small Scale Size Ratio (Feedstock) 25 7,000 5.3 

Large/Small Scale Cost Ratio 9.4 1,000 3.8 

Size Ratio/Cost Ratio 2.7 7 1.4 

Predicted Small Cost, Power Basis ($MM) 20.2 0.75 19.3 

Actual Small Cost, Design Basis ($MM) 20.5 0.39 16.1 

Percent Difference 1.5% 92.3% 19.8% 

Using a power basis provided a cost estimate well-within an order of magnitude acceptability for each 

technology, and, for some, within a level of acceptability (+/- 30%) for this level of detail. Although 

applying a power relationship to SMR reforming technology had the worst predictive capacity, the 

resulting estimate is still very close on an absolute basis. It was expected that this technology would have 

the most prediction difficulty because of the considerable difference in size relative to the other 

technologies chosen. 

A few general conclusions can be made regarding the cost relationship as technologies are scaled: 

 The power relationship that should be used depends largely on the size range that is being 

considered. For example, a very large difference in sizes (SMR technology) will usually be 

much more difficult to accurately predict than smaller differences. 

 Economies of scale are usually greater when technologies contain a novel or expensive main 

piece of equipment, such as the catalytic furnace in an SMR unit or a reactor/regenerator unit 

Task 1 Final Report ES-2 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
in a FCC design. Established technologies, such as refrigeration and fractionation in the NG 

Expander designs, are more easily predicted as to their impact when sizes change. 

 Many technologies have a breakpoint where the type of design changes due to size factors. 

This is true for all technologies considered in this study, as well as for many other 

technologies such as oxygen production units. As designs get close to this technology 

breakpoint, the capital cost per unit of feed changes considerably. 

Task 1 Final Report ES-3 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Introduction and Methodology 

This study evaluated chemical process designs that have been successfully scaled for large and small 

scale commercial applications. The goal was to provide additional insight into how the cost and design of 

mature commercial technologies scale as the overall unit throughput changes. To accomplish this, a list 

of process technologies was paired down to three for further evaluation: Fluidized Catalytic Cracking of 

refinery gas oil, Steam Methane Reforming of natural gas to hydrogen, and Natural Gas Liquids 

Expansion to separate the major components of natural gas. A list and description of all technologies 

considered for evaluation can be seen in Appendix A. 

Complete designs for each technology on a large and small basis included process flow diagrams, 

equipment lists, stream properties, operating costs, and installation factors. After the designs were 

completed, comparisons were made between the different scales to determine the cost impacts. Finally, 

conclusions were drawn for how general scaling approximations can be used regardless of the technology. 

The methodology for designing large and small schemes for each process was somewhat different. For 

the natural gas reforming and expansion technologies, the processes involved are non-proprietary. The 

equipment involved is standard to any chemical process industry: pumps, heat exchangers, distillation 

columns, pressure swing absorbers, and compressors. Based on the plant sizes and the range of plant 

information desired by NREL, complete theoretical facilities were designed by the project team. 

Information on all process flows, equipment sizes, energy requirements, and plant costs are presented in 

this deliverable. 

FCC units, however, are a proprietary technology with specific detailed information unavailable to the 

project team. A number of licensers (UOP, ExxonMobil, KBR, and Stone and Webster) were contacted 

to obtain additional information about designs of FCC units at both large and small scales. This 

information was supplemented from literature sources and previous FCC studies worked on by a major 

EPC vendor. Per NREL guidelines, the area of the FCC unit focused on by the project team is the 

reactor/regenerator section. Process design information for this area of the plant was obtained from both 

licensers and literature sources for units already constructed at large and small scales. In addition, scaling 

information and cost curves representing a broad range of sizes have been obtained to check the results. 

Typical stream conditions for both units have been estimated from vendor information and literature data. 

Finally, ICARUS model runs, using vendor and literature sizes, were run to confirm information supplied 

by other sources for plant costs. 

Task 1 Final Report 1 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Section 1 Fluidized Catalytic Cracking 

1.1 INTRODUCTION AND METHODOLOGY 

Data was collected from publications, licensors, previous in-house designs, and owner/operators of Fluid 

Catalytic Cracking (FCC) units. The focus of this research was on the capital and operating costs of large 

and small scale FCC reactor/regenerators. The product recovery, gas separation, and flue gas treating 

sections are not included. Since this study includes both large and small scale designs, two process flow 

diagrams (PFDs) have been developed to reflect the type of configuration which is considered most 

economic for the unit size. This has led to two different designs: 

 A “side-by-side” configuration of the reactor and regenerator, generally applicable to large 

scale units processing 20,000 or more barrels per stream day of fresh feed. This was used for 

the large scale unit, with a feed rate of 62,000 barrels per stream day (BPSD). 

 A “stacked” configuration of the reactor and regenerator, generally applicable to small scale 

units processing less than 20,000 BPSD. This was used for the small scale unit, with a feed 

rate of 2,500 BPSD. 

The equipment shown in these diagrams are based on the published designs of UOP. The costs 

established using the UOP design is representative of typical FCCs. All the major licensors confirmed 

that units of the same capacity designed by various licensors have similar capital costs. The variations are 

within the accuracy of this study of +/- 30%. 

Design and cost information was obtained from three main sources: 

 Direct Industry Contact: 

− Information for the large scale FCC design is based off a detailed design study performed 

by a major EPC vendor in the early 1990’s. 

− Design information for portions of the small scale FCC comes directly from UOP data on 

an actual facility. 

 Literature Sources: 

− A number of sources, footnoted in this report, were used for estimating operating costs 

and for providing checks on the industry derived numbers. 

 Commercially Available Software: 

− ICARUS was used as a check against industry supplied numbers, and to provide 

estimates for pieces of equipment not costed by other sources. 

Task 1 Final Report 1-1 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
1.2 TECHNOLOGY BACKGROUND 

FCC technology was developed in the late 1930s and early 1940s due to the demand for aviation fuel in 

World War II. The first commercial FCC unit was brought on-stream in May 1942 at what is now known 

as the ExxonMobil Refinery in Baton Rouge, Louisiana. The FCC process converts gas oil and heavier 

streams in a refinery to lighter, more valuable products via high temperature catalytic cracking. Typical 

FCC products are propylene, butylenes, gasoline, coke, and dry gas (light ends, typically saturated 

hydrocarbons in the pentane and lighter range). 

This technology is currently licensed by several companies. The main licensors include UOP LLC, EOM 

(ExxonMobil) / KBR (Kellogg Brown and Root) Alliance, and Stone and Webster Inc., each offering 

designs and features that enhance performance, reliability and flexibility of their units to accommodate a 

variety of feedstock and desired products. 

There are more than 400 FCC units operating worldwide, from 2,000 to 135,000 BPSD of fresh feed. All 

the FCC units recently built are outside of the United States, with the largest being the Reliance FCC in 

Jamnagar, Gujarat, India. This unit was designed by UOP with an original nameplate design capacity of 

135,000 BPSD. UOP reported that this unit achieved a record-breaking throughput of 180,000 BPSD in 

February 20031. In the United States, the capacity of FCC units range from 2,400 to 110,000 BPSD of 

fresh feed. The smallest FCC unit is at Montana Refining in Great Falls, Montana. This unit was 

designed for 2,000 BPSD by UOP and built in the 1950s. The small-scale design in this report reflects 

process conditions for a unit of this size. Very few facilities in the US exist under 10,000 BPSD due to 

weak process economics; many facilities that existed in this range have either shutdown or expanded. 

Over the years, with improvements in technology and increasing gasoline demand, most of the refiners 

have modified their FCC units to increase capacity. Several FCC units have been revamped to process 

more than 100,000 BPSD. According to the major licensors, the typical capacity of new grassroots FCC 

units in the United States is 40,000 to 60,000 BPSD. The large-scale design reflects the high end of this 

capacity, 60,000 BPSD. 

Operating conditions of FCCs vary depending upon the feed, desired products, and mechanical 

limitations. The typical ranges of operational parameters are provided in a number of literature sources2

. 

In addition, at the FCC Forum in May 2000, Engelhard (a major FCC catalyst manufacturer) presented 

the results of a FCC Benchmark study3

. The study concluded that following are typical operating 

conditions of an “average” FCC in the United States: 

 Processes 49,697 BPD of gas oil feed at an API gravity of 24.3 with 1,037 ppm nitrogen, 0.79 

wt % sulfur, 0.9 wt % carbon residue, and 174o

F aniline point. 

 Operates at a reactor temperature of 978°F and a regenerator temperature of 1,299°F. 

 Circulates 35.1 tons per minute of catalyst or 1.01 tons catalyst per barrel of feed. 

 Uses catalyst at a rate of 6.2 tons per day or 0.25 pounds per barrel of feed. 

 Produces 4.96 wt % coke and 3.23 wt % dry gas at a conversion level of 76 vol %. 

 

1 “Reliance FCC Unit Sets Capacity Record”, 4 March 2003, UOP Press Release, available at 

http://www.uop.com/objects/PR_RelCapacity.pdf 2 The sources found best for this study were “Petroleum Refining—Technology and Economics” by J. H. Gary and G. E. Handwerk and

“Handbook of Petroleum Refining Processes” by R.A. Meyers. 3 Information presented by The FCC Network, Volume 4, June-July 2000, available at 

http://thefccnetwork.com/pdf/newsletters/newsletter4.pdf 

Task 1 Final Report 1-2 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
 Yields 8.32 vol % propylene, 8.50 vol % butylenes, and 58.26 vol % gasoline with octane of 

92.3 RON and 80.8 MON. 

1.3 PROCESS DESCRIPTIONS AND DESIGN ASSUMPTIONS 

The FCC process uses a catalyst in the form of very fine particles, averaging about 70 microns in size, 

which behave as a fluid when aerated with a vapor. The fluidized catalyst is circulated continuously 

between the reactor and the regenerator, and acts as a vehicle to transfer heat from the combustion of coke 

in the regenerator for heating the feed and for the endothermic cracking reaction. Until about 1965, most 

units were designed with a discrete dense-phase fluidized catalyst bed in the reactor vessel. The units 

were operated so most of the cracking occurred in the reactor bed. In the 1960’s, the development of 

more reactive zeolite catalysts was a major breakthrough in catalyst technology and led to process design 

changes to obtain high selectivity and control of the cracking reactions in the “riser” or transfer line. The 

reactor vessel serves as the disengaging space for the product vapor and the spent catalyst, as well as the 

housing for the cyclones. 

The cracking reaction occurs in a vertical reactor riser in which a preheated liquid oil stream contacts hot 

catalyst from the regenerator. The oil vaporizes and cracks to lighter products as it moves up the riser and 

carries the catalyst powder along with it. Coke, a carbonaceous material having a low ratio of hydrogen 

to carbon (H/C) is a byproduct which deposits on the catalyst and reduces the catalyst activity. The spent 

catalyst and the converted products are separated in the reactor vessel. The cracked product vapors are 

sent to a fractionator and gas separation unit. Some of the hydrocarbons adsorbed on the spent catalyst 

are removed by steam stripping at the bottom of the reactor vessel. The spent catalyst passes to a separate 

vessel, the regenerator, where the coke is combusted to regenerate the catalyst. The hot, regenerated 

catalyst then passes to the bottom of the reactor riser, providing the heat needed in the reaction section. 

Two basic types of FCC units in use today are the “side-by-side” type and the Orthflow or “stacked” type. 

The reactor and regenerator are separate vessels adjacent to each other for the “side-by-side” type, which 

is typical of designs by UOP, ExxonMobil, and Stone and Webster. The reactor is mounted on top of the 

regenerator for the “stacked” type, which is typical of designs by KBR. 

Flow diagrams for both large and small scale cases can be seen in Figures 1-1 and 1-2, with 

corresponding stream properties in Tables 1-1 and 1-2. The design and cost estimates for each case are 

presented in the equipment list, which can be seen in Tables 1-3 and 1-4. The equipment list groups 

process equipment by the following categories: reactors, cyclones, columns/vessels/ tanks, and 

compressors. 

Task 1 Final Report 1-3 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
TABLE 1-1 LARGE SCALE FCC STREAM PROPERTIES 

PFD Stream 

No. Description Large Scale FCC 

1 Feed BPSD 62,000 

2 Product Vapor MMSCFD 100 

3 Combustion Air MMSCFD 146 

4 Catalyst Circulation Tons/min 50 

5 Regenerator Flue Gas MMSCFD 146 

FIGURE 1-1 LARGE SCALE FCC DESIGN 

Task 1 Final Report 1-4 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Task 1 Final Report 1-5 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 

TABLE 1-2 SMALL SCALE FCC STREAM PROPERTIES 

PFD Stream 

No. Description Small Scale FCC 

1 Feed BPSD 2,500 

2 Product Vapor MMSCFD 4 

3 Combustion Air MMSCFD 6 

4 Catalyst Circulation Tons/min 2 

5 Regenerator Flue Gas MMSCFD 6 

FIGURE 1-2 SMALL SCALE FCC DESIGN
TABLE 1-3 LARGE SCALE FCC EQUIPMENT LIST 

Design Operating 

Quantity Price, total Total Installed 

Item No Description Type Per Train Size Weight P T P T Power Useage Materials (unistalled) Cost 

lbs PSIG oF PSIG oF (No.) HP 

2002 Cost Index 

$US (US $) 

Reactors 

Reactor Vessel Fluidized Bed 1 27'-10"ID x 25'-0"T/T 1,300,000 50 1100 35 1000 

CS+1/16" CA (Shell), 

4" refractory $9,958,080 

--Stripper (Part of Reactor) Fluidized Bed 1 13'-7"ID (bottom of reactor) incl. w/ reactor 50 1100 35 1000 304SS+1/16" CA, 4" refract. incl. with reactor 

CS+1/16" CA (Shell), 

Regenerator Vessel Fluidized Bed 1 43'-8"ID x 35'-0" T/T incl. w/ reactor 50 1600 35 1300 4" refractory incl. with reactor 

--Combustor (Part of Regenerator) Fluidized Bed 1 38'-0"ID x 15'-0" T/T incl. w/ reactor 50 1600 35 1300 304SS+1/16" CA, 4" refract. incl. with reactor 

Total $9,958,080 Cyclones 

10 FT2 inlet, 11 FT2 outlet, 2.5' 

Reactor Cyclones Cyclone 4 dipleg, 2 stage 42,400/unit 50 1100 35 1000 1-1/4 Cr -1/2 Mo + 1/8" CA $910,200 

5.5 FT2 inlet, 6.3 FT2 outlet, 2' 

Regenerator Cyclones Cyclone 10 dipleg, 2 stage 35,000/unit 50 1600 35 1300 304SS + 1/16" CA $2,484,600 

$3,394,800 

Total Columns, Vessels & Tanks 

Riser Vertical 1 5'-9"ID x 98'-0" T/T incl. w/ reactor 50 1600 35 1300 CS+1/16" CA, 4" refract. incl. with reactor 

Fresh Catalyst Storage Drum Vertical 1 15'-6"ID x 62'-0" T/T 112,000 50 650 50 650 CS+1/16" CA $214,020 

Spent Catalyst Storage Drum Vertical 1 15'-6"ID x 62'-0" T/T 112,000 50 650 50 650 CS+1/16" CA $214,020 

Total $428,040 Compressors & Blowers 

Combustion Air Compressor Centrifugal 1 171000 ACFM 20,300 CS $8,085,000 

Total $8,085,000 Other 

Reactor Vent Silencer Vertical 1 4'-6"ID incl. w/ reactor 50 1100 35 1000 CS+1/16" CA $41,820 

Total $41,820 

SUBTOTAL $21,907,740 

BULK MATERIALS $28,480,062 N

TOTAL PROJECT COST $50,387,802 $102,615,654 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
TABLE 1-4 SMALL SCALE FCC EQUIPMENT LIST 

Design Operating 

Quantity Price, total Total Installed 

Item No Description Type Per Train Size Weight P T P T Power Useage Materials (unistalled) Cost 

lbs PSIG oF PSIG oF (No.) HP 

2002 Cost Index 

$US (US $) 

Reactors 

Reactor Vessel Fluidized Bed 1 7' ID x 33' T/T 100,000 50 1100 35 1000 

CS+1/16" CA (Shell), 

4" refractory $1,550,000 

U

--Stripper (Part of Reactor) Fluidized Bed 1 3'-6"ID (bottom of reactor) incl. w/ reactor 50 1100 35 1000 304SS+1/16" CA, 4" refract. incl. with reactor 

CS+1/16" CA (Shell), U

Regenerator Vessel Fluidized Bed 1 9"-6" ID x 40'-0" T/T incl. w/ reactor 50 1600 35 1300 4" refractory incl. with reactor 

--Combustor (Part of Regenerator) Fluidized Bed 1 8'-0"ID x 6'-6" T/T incl. w/ reactor 50 1600 35 1300 304SS+1/16" CA, 4" refract. incl. with reactor 

Total $1,550,000 Cyclones 

Reactor Cyclones Cyclone 2 2 foot diameter, single stage 3,100/unit 50 1100 35 1000 1-1/4 Cr -1/2 Mo + 1/8" CA $30,000 

Regenerator Cyclones Cyclone 3 3 foot diameter, two stage 5,200/unit 50 1600 35 1300 304SS + 1/16" CA $90,000 

$120,000 

Total Columns, Vessels & Tanks 

Riser Vertical 1 2' ID x 85'-0" T/T incl. w/ reactor 50 1600 35 1300 CS+1/16" CA, 4" refract. incl. with reactor Fresh Catalyst Storage Drum Vertical 1 3'-6"ID x 12'-0" T/T 20,000 50 650 50 650 CS+1/16" CA $36,900 

Spent Catalyst Storage Drum Vertical 1 3'-6"ID x 12'-0" T/T 20,000 50 650 50 650 CS+1/16" CA $36,900 

Total $73,800 Compressors & Blowers 

Combustion Air Compressor Centrifugal 1 7000 ACFM 1,600 CS $600,000 

Total $600,000 Other 

Reactor Vent Silencer Vertical 1 1' ID incl. w/ reactor 50 1100 35 1000 CS+1/16" CA $6,765 

Total $6,765 

SUBTOTAL $2,350,565 

BULK MATERIALS $3,055,735 

TOTAL PROJECT COST $5,406,300 $11,010,025 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Two cost factors were used on top of the equipment cost to come up with the total project 

investment (TPI). First, the equipment cost was multiplied by 1.3 to estimate the cost for all bulk 

materials. This includes items such as instruments, piping, concrete, electrical, and all associated 

labor involved with fabricating these items. An installation and project cost factor of 2.04 was 

applied to the equipment cost to arrive at the TPI. This not only includes installation, but also all 

indirect costs. Both factors were based upon Nexant and vendor estimates for a detailed FCC 

design. The results of each case can be seen in Table 1-5. 

TABLE 1-5 TOTAL FCC PROJECT COSTS, 2002 BASIS 

Large Scale Plant Small Scale Plant 

Feed Rate (BPSD) 62,000 2,500 

Reactor/Regen TPI ($MM) 102.6 11.0 

Total Project Investment ($MM) 191.6 20.5 

Total Investment ($/bbl) ~3,000 ~8,200 

As can be seen from Table 1-5, there is a significant economy of scale advantage realized in the 

construction of larger FCC units. In fact, the $/bbl investment for even larger plants (~100,000 

BPSD) may approach ~$20004. This is a large reason why small FCC units are no longer built, 

except in specialty niche applications where petroleum supply and product demand are limited 

but stable. In general, petroleum companies are building all refinery process units large to 

improve refining margins. 

1.3.1 Large Scale (Side-by-Side) FCC 

Riser Section 

In the “side-by-side” design, the riser is inside the steam stripper, as shown in Figure 1-1. The 

operating pressure ranges from 8 to 30 psig, and is generally limited to 15 to 20 psig by the 

design and pressure balance of the reactor and the regenerator. Lower operating pressure tends 

to improve yield selectivity, enhance stripper performance, and reduce air blower requirements, 

but increases vessel sizes and investment cost. 

The riser temperature ranges from 950o

F to 1025oF. It is controlled by the catalyst circulation 

rate, which in turn is controlled by the position of the slide valve at the bottom of the 

Regenerated Catalyst Standpipe. Increasing temperature, catalyst/oil ratio, catalyst activity, and 

contact time results in higher conversion and significant octane improvement. However, 

cracking at higher temperatures also produces more gas and coke which are less desirable 

products. The optimum temperature is when just enough coke is formed on the catalyst so that 

the heat of combustion in the regenerator can provide all the heat required for reaction when the 

catalyst returns to the riser-reactor. 

 

4 Information from Mark Schnaith, UOP FCC Business Development Manager, personnel communications, and “Refinery Processes 2002”, 

Hydrocarbon Processing, November 2002. The actual investment will depend on the technology chosen, the gas oil feed properties, and 

the products desired. 

Task 1 Final Report 1-8 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Since modern catalysts are highly active, conversion is controlled by varying catalyst circulation 

rate. A short residence time of the gas in the riser, usually less than 5 seconds, is maintained in 

order to produce the desired product selectivity and amount of coke formation. 

Fresh catalyst is added to maintain the desired activity of the catalyst as the activity slowly 

declines from residual coke deposit and deterioration. Spent catalyst is removed periodically to 

maintain the inventory in the desired range. 

Reactor Section 

The reactor serves as a disengaging vessel for the catalyst and product vapor, and as the housing 

for the cyclones. Operating pressure and temperature of the reactor are similar to those at the 

riser. 

At the top of the riser, all the selective cracking reactions are terminated by separation of the 

product vapor from the spent catalyst. Since over-cracking diminishes the value of the product, 

it is important to minimize the product vapor residence time in the reactor disengaging vessel. 

Special mechanical riser termination devices have been developed to ensure rapid separation of 

the vapor from the catalysts. These devices include vortex separation system developed by UOP, 

vented risers, and close-coupled cyclones often used by KBR/EOM. Cyclones mounted inside 

the reactor removes the catalysts from the product vapor before entering the fractionator and gas 

recovery sections. The diplegs of these cyclones are sealed in the fluidized catalyst bed or with 

counter weighted valves. 

Spent catalyst separated from the cyclones contains entrained and adsorbed hydrocarbons as well 

as the coke deposited by the cracking reaction. Removing the hydrocarbons from the catalyst 

before regeneration minimizes the regenerator bed temperature, regeneration air requirement, 

and flue gas volume. Spent catalyst flows down into the Spent Catalyst Stripper which is 

connected to the base of the reactor vessel. In this section, ascending stripping steam is used to 

remove as many hydrocarbons as possible from the catalyst as it descends through the stripper. 

Hydrocarbon vapors and steam flow upwards into the reactor disengaging vessel and mix with 

the product vapors from the riser. The stripped spent catalyst descends through the Spent 

Catalyst Standpipe into the regenerator. A refractory-lined steel slide valve at the base of this 

standpipe provides control of the catalyst circulation rate. 

In the “side-by-side” design, the reactor vessel is located next to the regenerator vessel. The 

Spent Catalyst Standpipe and the Regenerated Catalyst Standpipe connect these two vessels, as 

shown in Figure 1-1. 

Regenerator Section 

The regenerator is a fluidized bed where the activity of the spent catalyst is restored. Coke 

deposited on the catalyst during the cracking reactions is burnt off with air. Pressurized air is 

injected into the regenerator to provide oxygen for coke combustion and bed fluidization. The 

exothermic process of coke combustion heats the catalyst which transports the heat energy to the 

riser for the cracking reactions. Cyclones mounted inside the regenerator vessel remove the 

catalyst carryover from the flue gas before it is cooled and treated for release to the atmosphere. 

Task 1 Final Report 1-9 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Both small and large scale units have a combustor style regenerator. This design features a fast 

fluidized bed as the low inventory carbon burning zone followed by a higher velocity transportriser heat exchange zone. It provides uniform coke-air distribution which enhances the ability to 

burn almost all the coke to carbon dioxide, leaving very low levels of carbon monoxide in the 

flue gas. The regenerated catalyst has low residual carbon and high effective activity, and is 

returned to the riser through the Spent Catalyst Standpipe. 

Typical regenerator operating pressure ranges from 15 to 30 psig. The regenerators are generally 

operated at the maximum practical temperature from 1,200 to 1,500o

F for maximum combustion 

of the coke deposits and preservation of catalyst activity and selectivity. The higher the 

differential temperature between the regenerator and the reactor, the lower the catalyst 

circulation needed. 

Design Assumptions 

Since the methodology for design of an FCC of this scale is well-established, a high level of 

confidence exists in the estimates presented here. Although this design was originally performed 

in the early 1990’s, the major design aspects have not changed. Few assumptions were required 

due to the quality of the existing data. Assumptions used in the design include: 

 Cost Escalation: Factors were used to escalate the early 1990’s cost data to 2002. 

These factors largely came from the ICARUS database. Not all items escalated at the 

same rate; for example, the escalation on compressor costs was considerably less than 

on the reactor/regenerator vessels. 

 Scope: Per NREL guidelines, an entire FCC design is not presented here. The items 

deemed to be most important for future NREL use are the reactor/regenerator and 

associated equipment. Since all vendor supplied FCC design packages are for an 

entire FCC facility including downstream fractionation and product treating, splitting 

the design to just focus on the reactor/regenerator made it difficult to estimate utility 

requirements. Main process utilities were therefore estimated from literature 

estimates instead of vendor packages. 

 Internals: Items such as the combustor, riser, cyclones, and stripper are usually 

provided as a package in the FCC design estimate, with limited design information 

provided. Some of these pieces of equipment are considered proprietary by vendors. 

1.3.2 Small Scale (Stacked) FCC 

In general, all the major licensors agree that the “stacked” design requires less plot space and less 

structural steel, thus lowering the installed cost. As capacity increases, the vessel sizes also 

increase to a point where structural cost exceeds this advantage. As a rule of thumb, a capacity 

of 20,000 BPSD is the cut-off size where “side-by-side” design becomes more economical than 

“stacked” design, according to UOP. There are other factors that affect the cost/BPSD for the 

same capacity and thus affect the selection of the FCC configuration. These factors include 

location, available land space, and local building codes. 

In general, the theory behind the design of the small-scale unit is very similar to that of the largescale unit. Regardless of the variation of design by the licensors, the components that make up 

Task 1 Final Report 1-10 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
the FCC are the same: the riser, the reactor, and the regenerator. Two design differences can be 

seen in the riser and reactor for the stacked design: 

 The riser is outside the stack of reactor, stripper and regenerator vessels. The product 

vapors and catalyst flow from the riser into the reactor vessel through a nozzle. 

 The reactor vessel is mounted on top of the regenerator vessel. The Spent Catalyst 

Standpipe is located inside the regenerator vessel. The Regenerated Catalyst 

Standpipe connects the regenerator to the external riser. 

Design Assumptions 

Less vendor detail and previous design information was available for estimation of the small 

scale FCC. In general, small scale FCCs are no longer economic, and very few units under 

20,000 BPSD are being built today. UOP was able to provide some design information for its 

2,500 BPSD unit in Great Falls, Montana, which was built in the 1950’s. A greater level of 

assumptions and scaling was required when developing this design estimate: 

 Reactor/Regenerator: UOP provided the sizes of both the reactor and regenerator, but 

not the cost. The cost was developed through the use of ICARUS and Nexant design 

information. Internals were estimated from literature sources. 

 Cyclones: The quantity, dimensions, and cost for the cyclones were developed 

through the use of ICARUS and vendor input. 

 Air Compressor: Literature factors relating the amount of process air to the feedrate 

was used as the basis for the air compressor sizing. Horsepower requirements and 

costs were obtained through the use of HYSYS and ICARUS. 

Task 1 Final Report 1-11 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
1.4 OPERATING COSTS AND UTILITY REQUIREMENTS 

The basis for the operating costs and utility requirements are shown in Table 1-6. 

TABLE 1-6 FCC OPERATING COSTS AND UTILITY REQUIREMENTS 

Operating Variable Large Scale Small Scale Notes 

Catalyst5

 Initial Load 

 Make-Up

468 tons 

6 tons/day 

19.5 tons 

0.25 tons/day 

$2000/ton catalyst cost6. 

Assumes no enhanced olefins 

production 

Combustion Air 146 MMSCFD 6 MMSCFD Assumes no excess air 

Medium Pressure Steam Consumption 

 (React/Regen Only)7 68,333 lb/hr 1,424 lb/hr Side-by-side plant requires 

additional steam 

High Pressure Steam Production 

(Overall FCC)8 100,000 lb/hr 4,167 lb/hr Overall FCC is net steam long 

Electricity Use9 15 MW 0.6 MW Assumes electric driver for air 

compressor 

The majority of the estimates used here come from literature sources. When sources differed in 

their estimates, an average was assumed to develop a “typical” FCC operating cost. For the most 

part, FCC utility requirements can be fairly accurately predicted just by knowing the unit feed 

rate: catalyst, combustion air, electricity, and steam requirements can all be roughly estimated by 

knowing this key quantity. A few key items for consideration when developing overall operating 

costs: 

 Daily catalyst make-up requirements can vary considerably based on the age of the 

unit, the outlet composition desired, and the efficiency of the cyclones. 

 Most FCC units use some excess process air to assure complete combustion of the 

coke deposited on the catalyst. Since this can vary considerably, a base design 

assumption without process air is used for both units. The design for the air 

compressors assumes capacity available for excess process air. 

 Steam costs are difficult to estimate, for an FCC will be integrated into an overall 

refinery configuration. It is likely that sufficient process steam will be available to 

keep the cost of this utility low. 

 

5 Data compiled from the following sources: “Refinery Processes 2002”, Hydrocarbon Processing, November 2002, “Petroleum Refining—

Technology and Economics” by J. H. Gary and G. E. Handwerk, 1984, and The FCC Network, Volume 4, June-July 2000, available at 

http://thefccnetwork.com/pdf/newsletters/newsletter4.pdf 6 Information from Mark Schnaith, UOP FCC Business Development Manager, personnel communications, 15 August 2005. 

7 Data gathered from Bechtel survey of FCC utility requirements, 1981. Average of UOP designs taken, with distinctions made between 

stacked and side-by-side design requirements. 8 Information from KBR data in Hydrocarbon Processing, November 2002, and UOP information from Meyers R., “Handbook of Petroleum

Refining Processes”, Third Edition, 2004. 9 Bechtel survey of UOP equipment, 1981. Data checked versus expected compressor horsepower requirements. 

Task 1 Final Report 1-12 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Section 2 Natural Gas Reforming 

2.1 INTRODUCTION AND METHODOLOGY 

Detailed designs for both large and small scale Steam Methane Reforming (SMR) units were 

developed specifically for hydrogen production. The small scale design assumes a methane feed 

rate of 0.024 MMSCFD (producing 10 lb/hr of hydrogen), while the large scale design uses a 

methane feed rate of 168 MMSCFD (producing ~100,000 lb/hr of hydrogen). These feed rates 

represent a typical size for two distinct types of SMR facilities: the small scale plant is consistent 

with hydrogen production for small fuel cell applications, while the large facility would be seen 

at a petrochemical or ammonia plant. 

Data was collected from publications and licensors of SMR units. The focus of this research was 

on the capital and operating costs for entire large and small scale units including natural gas 

purification, reforming, shift, and hydrogen purification. The compression and storage of the 

hydrogen product are not included. Design and cost information was obtained primarily from 

literature sources and process modeling. A number of sources, noted in this report, were used for 

estimating capital and operating costs. 

2.2 TECHNOLOGY BACKGROUND 

Natural gas reforming is the process of converting natural gas feedstocks to a synthesis gas 

consisting of CO, CO2, and H2 using steam as a reactant. The synthesis gas will also contain 

small amounts of unreacted methane and residual steam. This process is the first step in many 

industrial processes including methanol synthesis, ammonia manufacture, gas-to-liquids (GTL), 

and hydrogen production. The earliest SMR units were introduced in the U.S. in the 1930s by 

Standard Oil of New Jersey at their Bayway and Baton Rouge refineries for hydrogen production 

from offgases10. 

Steam methane reforming is the most common and least expensive route to producing hydrogen. 

The technology is well developed and commercially available at a wide capacity range, from less 

than 1 MMSCFD of methane feed for fuel cells, to 400 MMSCFD of methane feed for ammonia 

production. The process involves the catalytic reaction between methane and steam to produce a 

synthesis gas. The synthesis gas is further processed through another catalytic step to increase 

the hydrogen fraction and decrease the carbon monoxide via the water gas shift reaction. 

Finally, a purification step removes all other components to produce the hydrogen product. 

2.3 PROCESS DESCRIPTIONS AND DESIGN ASSUMPTIONS 

Similar to the information presented in the FCC section, flow diagrams, stream properties, and 

an equipment list for both cases is presented below. Due to the larger overall process scope and 

greater process complexity, detailed stream properties are shown for each case. 

Unlike the FCC design case, unique installation factors specific to SMR units have not been used 

or identified by the Nexant team. Therefore, a process similar to that used by NREL in the 

recent “Biomass Syngas to Hydrogen Production Design Report” was used. After calculation of 

 

10 Catalyst Handbook, Second Edition, Martyn V. Twigg, ed., Wolfe Publishing Ltd., 1989 

Task 1 Final Report 2-1 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Task 1 Final Report 2-2 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 

the equipment cost, a factor of 2.47 was used to calculate the Total Installed Cost (TIC). For 

indirect costs, factors developed by NREL to account for engineering costs, construction, 

contingencies, and fees were applied. Indirect cost factors were then added to the TIC to obtain 

the Total Project Investment (TPI). A summary of the results for both cases can be seen in Table 

2-1. 

TABLE 2-1 SMR PROJECT COSTS, 2005 BASIS 

Large Scale Plant Small Scale Plant 

Natural Gas Feed Rate (MMSCFD) 167.8 0.024 

Hydrogen Production Rate (lb/hr) 101,885 10.6 

Equipment Cost ($MM) 98.8 0.10 

Installed Cost ($MM) 244.0 0.26 

Total Plant Investment ($MM) 368.5 0.39 

Total Investment/SCFD Feed ($) 2.2 16.2 

Due to the very different plant sizes being considered, a large economy of scale difference is 

evident. Roughly 7 times the investment per SCFD of methane feed is required to produce small 

amounts of hydrogen relative to the largest plants in operation today. While the capital required 

is significant to capture this economy of scale, large scale centralization of hydrogen production 

appears to have a significant cost advantage if large amounts of hydrogen is desired. 

2.3.1 Small Scale SMR Design

Figure 2-1 shows the PFD for the small scale SMR unit, while Table 2-2 shows the stream

conditions. The equipment list for the small scale design can be found in Table 2-3. 
Natural Gas Compressor HDS Preheater

HDS/ZnO Parallel Beds

Steam Methane

Reforming

Natural Gas

1

2

Water

4 Water Pump

Steam 

Boiler

3

5

Reformer Cooler

High Temperature Shift

6

Condenser

7 Fuel G

Water Knockout

8

Water

Hydrogen Purification

9

Hydrogen Product

Natural Gas Burner

Burner Exhaust

FIGURE 2-1 SMALL-SCALE SMR PROCESS FLOW DIAGRAM 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
TABLE 2-2 SMALL SCALE SMR STREAM PROPERTIES 

Stream ID Natural Gas Feed 

Desulfurized 

Natural Gas Reaction Steam Reformer Feed 

Stream Number 1 2 3 4 Methane (lb-mole/h) 2.13 2.13 2.13 H2O 0.00 0.00 7.36 7.36 CO CO2 0.02 0.02 0.02 Hydrogen Ethane 0.08 0.08 0.08 Propane 0.04 0.04 0.04 Oxygen 0.02 0.02 0.02 Nitrogen 0.04 0.04 0.04 Total lbmole/h 2.4 2.4 7.4 9.7 Mass Flow [lb/hr] 41.7 41.7 132.6 174.3 Std Gas Flow [cu ft/hr] 892 892 2,793 3,685 Std Gas Flow [MMSCFD] 0.02 0.02 0.07 0.09 Temperature [F] 850 850 850 847 Pressure [psia] 295 295 295 295 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
TABLE 2-2 SMALL SCALE SMR STREAM PROPERTIES (Cont’d) 

Stream ID Shift Feed Shifted Gas 

Cooled Shifted 

Gas 

Shift 

Condensate PSA Feed 

Hydr

Pro

Stream Number 6 7 8 9 1Methane (lb-mole/h) 0.63 0.63 0.63 0.00 0.63 0.H2O 4.76 3.79 3.79 3.70 0.08 0.CO 1.01 0.04 0.04 0.00 0.04 0.CO2 0.82 1.79 1.79 0.00 1.79 0.Hydrogen 6.04 7.01 7.01 0.00 7.01 5.Ethane 0.00 0.00 0.00 0.00 0.00 0.Propane Oxygen 0.02 0.02 0.02 0.00 0.02 0.Nitrogen 0.04 0.04 0.04 0.00 0.04 0.Total lbmole/h 13.3 13.3 13.32 3.7 9.6 5Mass Flow [lb/hr] 174.3 174.3 174.3 66.8 107.4 10Std Gas Flow [cu ft/hr] 5,056 5,056 5,056 1,406 3,650 1,9Std Gas Flow [MMSCFD] 0.12 0.12 0.12 0.03 0.09 0.Temperature [F] 840 480 130 130 130 13Pressure [psia] 290 290 285 280 275 42

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
TABLE 2-3 SMALL-SCALE SMR EQUIPMENT LIST 

Design Operating 

Quantity Price, total Total Installed 

Item No Description Type Per Train Size Weight P T P T Power Useage Materials (unistalled) 

2002 Cost Index 

Cost 

lbs PSIG oF PSIG oF HP $US (US $) 

Reactors 

$5,277 

Hydro-Desulfurization Bed Fixed Bed 2 4" pipe, 30" length 

Shell: 14" D, 40" L 

350 1,000 295 850 316 SS 

Incoloy 800H (Shell), 

Reforming Bed Fixed Bed 1 Tubes: 134 @ 1/2" D, 39" L 350 1,600 295 1,475 Haynes 556 (Tubes) $28,746 

Shell: 15.25" D, 75" L 

High Temperature Shift Fixed Bed 1 Tubes: 106 @ 1" D, 75" L 350 1,000 295 840 316 SS $20,594 

Pressure Swing Adsorber Fixed Bed 4 12" D, 72" L 350 150 295 130 316 SS $18,075 

Total $72,692 Exchangers 

Shell: 2.5" Sched 40, 12" long Included with HDS 

HDS Preheater Shell/Tube 1 Tube: 1" D, 23" L (1 tube) 350 1,000 295 850 316 SS, Incoloy Bed 

Shell: 16.5" D, 320" L 

Steam Boiler Shell/Tube 1 Tubes: 64 1" tubes 350 1,000 295 850 316 SS (shell), Incoloy (tubes) $18,415 

316 SS (shell), Included with HT Shift 

Reformer Cooler Air Fan 1 13 Tubes, 10"D, 18" L 350 1,600 295 1,475 Haynes 556 (tubes) reactor 

3.5 cu. ft, stainless steel finned 

Post-Shift Condenser Air Fan 1 tubes 350 600 295 480 316 SS finned tubes $4,729 

Total $23,144 Pumps, Compressors & Blowers 

Natural Gas Compressor Centrifugal 1 15 ACFM 350 500 295 200 8.5 HP CS $3,354 

Total 

Water Pump Pos. Displacement 1 1.5 L/min, 300 psig 350 200 200 300 0.06 HP CS $3,210 

$6,564 

SUBTOTAL $102,400 

TOTAL PROJECT COST $252,928 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Natural gas (stream 1) is compressed from near atmospheric delivery pressure to 295 psia to 

decrease downstream equipment sizes and take advantage of the increased reaction rates and 

driving forces seen at higher pressures. The gas is then heated in the hydrodesulfurization (HDS) 

preheater to 750°F before entering the HDS/ZnO desulfurization beds. This step is necessary 

because the reforming catalysts are poisoned by even trace quantities of hydrogen sulfide, 

mercaptans, and odorants. The HDS catalyst converts mercaptans to hydrogen sulfide while the 

zinc oxide adsorbs the hydrogen sulfide. Two parallel beds are used to allow undisturbed 

operations. Once the bed material is fully loaded, the bed can be emptied and refilled while the 

other bed is treating the gas. 

Steam (stream 3) is produced by vaporizing water in a boiler heated by a natural gas burner. 

Steam is necessary for heat during the endothermic reforming reaction and for use in the watergas shift step. The steam and hot desulfurized methane stream (stream 4) are combined and fed 

to the steam methane reformer. This reactor contains a nickel-based catalyst where the following 

reactions take place: 

CH4 + H2O Ù CO + 3H2 (reforming) 

CO + H2O Ù CO2 + H2 (water-gas shift) 

The reforming reaction is highly endothermic, with the heat of reaction supplied by an external 

fired heater. The steam to carbon ratio typically varies from 2.5 to 5 on a molar basis. The 

equilibrium conversion of methane at 295 psig is 73%. A depiction of a typical small scale SMR 

reactor can be seen in Figure 2-2. Designs follow fairly closely with standard heat exchanger 

design to minimize unique requirements that lead to increased cost. 

FIGURE 2-2 TYPICAL SMALL SCALE SMR REACTOR DESIGN11

The resulting synthesis gas is cooled through the HDS preheater and reformer cooler to enter the 

high temperature shift reactor (stream 6). Here the water gas shift reaction occurs over a 

promoted iron oxide catalyst to increase the amount of hydrogen while decreasing the carbon 

monoxide content. Over 95% conversion of the carbon monoxide is achieved. 

 

11 “Cost and Performance Comparison of Stationary Hydrogen Fueling Appliances, Task 2 Report” Myers, D.B., et al., Directed 

Technologies, Inc. (April 2002) 

Task 1 Final Report 2-7 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
The next step is to cool the gas stream in order to condense the water vapor in the water 

knockout vessel. Pressure swing adsorption (PSA) is then used to produce the final hydrogen 

stream (stream 10) at 99.99% purity and with less than 10 ppm carbon monoxide. The remaining 

hydrogen, methane, carbon dioxide, carbon monoxide, nitrogen, and other trace components are 

recovered in the regeneration step and used as a fuel gas (stream 11). The PSA system uses a 

number of adsorbent beds, containing a 5Å molecular sieve to selectively pass hydrogen while 

adsorbing the other components. Once loaded with these components and before they 

breakthrough the bed into the hydrogen product, the bed is taken off-line for regeneration while a 

parallel bed is put on-line with no break in the continuity of gas flow. The pressure in the 

regenerating bed is dropped to drive off the adsorbed components. The cycle is run 

automatically on a timed schedule to avoid breakthrough. A four-bed system is used in this 

design. 

Design Assumptions 

Since this study includes both large and small scale designs, two PFDs have been developed to 

reflect the type of configuration which is considered most common and economic for the unit 

size. This has led to two different designs: 

 A reformer that integrates the external heat to drive the endothermic reforming 

reactions and convection heat exchange to provide heating and cooling of the process 

streams. This design is typically used for medium and large systems, with the basis 

used for the large scale unit design. 

 A stand-alone reformer with external heating provided by natural gas and nonintegrated heat exchange systems designed to minimize cost. This was used for the 

small scale unit. The type of heat integration and reformer design used in the large 

scale unit is not practical for hydrogen production at the small scale. Creating 

designs that exactly replicated the technology used in either case would not be 

indicative of true commercial applications. For example, air coolers for the reformer 

cooler and post-HTS condenser are considerably more practical in the small scale due 

to the low duty required. 

The small scale plant assumes that the design can be used anywhere hydrogen is required, and 

does not require integration into an industrial facility. Compressors and pumps are included for 

natural gas and water, with the assumption that neither of these feeds is available at the 

conditions required in the process. Fuel gas produced in the process can be used in the natural 

gas burner to help provide the reformer heat duty. 

2.3.2 Large Scale SMR Design 

Figure 2-3 shows the PFD for the large scale SMR unit, while Table 2-4 shows the stream 

conditions. The equipment list for the large scale design can be found in Table 2-5. 

The major difference between the small and large scale systems is the addition of an internal 

direct-fired heater to provide the necessary heat of reaction for the large volume of natural gas 

feed. At this size, the reformer reactor is usually integrated with the fired heater assembly in a 

packaged system. In addition, the large scale system uses additional heat recovery systems to 

Task 1 Final Report 2-8 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Task 1 Final Report 2-9 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 

utilize heat from the direct fired heaters. The heat integration increases the thermal efficiency of 

the plant relative to the small scale unit. 

The natural gas feed stream is heated in the reformer exhaust system (stream 1) and passes 

through the desulfurization vessels where H2S is adsorbed on a solid adsorbent. At least two 

beds are operated in parallel so that one unit can be recharged when the bed material is fully 

loaded without shutting down flow. 

The desulfurized gas (stream 2) is mixed with steam (stream 3) and superheated in the feed 

preheat coil. The feed mixture (stream 4) then passes through reforming catalyst-filled tubes in 

the reformer vessel. In the presence of the nickel-based catalyst, the hydrocarbons (mainly 

methane) react with the steam to produce a synthesis gas composed mainly of carbon monoxide 

and hydrogen. The endothermic heat of reaction is provided by controlled external firing in the 

reformer. The exhaust gas from the fired heater provides a bank of convection coil heating for: 

 Steam generation 

 Feed/steam preheating 

 Steam superheating 

 Feed preheating 

 Condensate preheating, and 

 Combustion air preheating 

The synthesis gas (stream 5) leaving the reformer is cooled by the process gas cooler/steam

generator. Gas is then fed to the high temperature shift reactor (stream 6). This catalyst is 

copper promoted iron-chromium, which converts carbon monoxide and water to additional 

hydrogen with a coproduct of carbon dioxide. The shift reactor effluent (stream 7) is cooled in 

the condensed water heater, a boiler feed water (BFW) preheater, and a gas trim cooler. The 

cooled stream (stream 8) flows to a condensate separator where the remaining condensate is 

separated and the gas (stream 9) is sent to a PSA Hydrogen Purification System. 

The PSA system is on an automatic cycle to minimize operational attention. It cycles several 

vessels between adsorption and regeneration cycles where (1) essentially all the gas components 

except hydrogen are adsorbed on a molecular sieve material, allowing the hydrogen to pass at 

pressure (stream 10), and then (2) desorption of the adsorbed components by reducing the 

pressure in the vessel. The PSA off-gas (stream 11) is sent to the reformer to provide some of 

the fuel required for reforming. 
Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 

FIGURE 2-3 LARGE SCALE SMR PROCESS FLOW DIAGRAM 
Task 1 Final Report 2-11 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 

TABLE 2-4 LARGE SCALE SMR STREAM PROPERTIES 

Stream ID 

Natural Gas 

Feed 

Desulfurized 

Natural Gas Reaction Steam Reformer Feed 

Reformed 

Product 

Stream Number 1 2 3 4 5 

Methane (lb-mole/h) 16,707.4 16,707.4 - 16,707.4 3,864.7 

H2O 2.1 2.1 49,166.0 49,168.1 29,142.8 

CO - - - - 10,413.9 

CO2 184.0 184.0 - 184.0 4,989.7 

Hydrogen - - - - 49,101.1 

Ethane 663.3 663.3 - 663.3 0.0 

Propane 350.1 350.1 - 350.1 0.0 

Oxygen 184.0 184.0 - 184.0 184.0 

Nitrogen 331.1 331.1 - 331.1 331.1 

Total lbmole/h 18,429.4 18,422.0 49,166.0 67,588.0 98,027.2 

 

Mass Flow [lb/hr] 326,967 326,716 885,730 1,212,447 1,212,459 

Std Gas Flow [MMSCF/hr] 7.0 7.0 18.7 25.6 37.2 

Std Gas Flow [MMSCFD] 167.8 167.8 447.8 615.6 892.8 

Temperature [F] 1000 1000 1000 996 1615 

Pressure [psia] 430 430 430 430 430 

Stream ID Shift Feed Shifted Gas 

Cooled 

Shifted 

Gas PSA Feed 

Hydrogen 

Product PSA Purge 

Stream Number 6 7 8 9 10 11 

Methane (lb-mole/h) 3,864.7 3,864.7 3,864.7 3,864.7 0.4 3,864.3 

H2O 29,142.8 19,520.7 19,520.7 492.1 0.0 492.0 

CO 10,413.9 791.8 791.8 791.8 0.1 791.7 

CO2 4,989.7 14,611.7 14,611.7 14,585.8 1.5 14,584.3 

Hydrogen 49,101.1 58,723.1 58,723.1 58,722.1 50,501.0 8,221.1 

Ethane 0.0 0.0 0.0 0.0 0.0 0.0 

Propane 0.0 - - - - - 

Oxygen 184.0 184.0 184.0 184.0 0.0 183.9 

Nitrogen 331.1 331.1 331.1 331.1 0.0 331.1 

Total lbmole/h 98,027.2 98,027.2 98,027.2 78,971.5 50,503.0 28,468.5 

Mass Flow [lb/hr] 1,212,459 1,212,456 1,212,456 868,507 101,885 766,622 

Std Gas Flow [MMSCF/hr] 37.2 37.2 37.2 30.0 19.2 10.8 

Std Gas Flow [MMSCFD] 892.8 892.8 892.8 719.2 460.0 259.3 

Temperature [F] 840 525 130 130 130 130 

Pressure [psia] 425 425 420 420 420 3 
TABLE 2-5 LARGE-SCALE SMR EQUIPMENT LIST 

Design Operating 

Quantity Price, total 

Item No Description Type Per Train Size Weight P T P T Power Useage Materials (unistalled) 

2002 Cost Index 

Total Installed Cost 

lbs PSIG oF PSIG oF (No.) HP $US (US $) 

Reactors 

Hydro-Desulfurization Bed Fixed Bed 8 16' 6" L, 8' 6" D 

26' 0" L, 13' 6" D 

347,684 600 1000 430 750 SS316 

Incoloy 800H (Shell), 

Reforming Bed Fixed Bed 1 1788 Tubes 546,360 600 1800 430 1650 Haynes 556 (Tubes) 

High Temperature Shift Fixed Bed 2 27' 0" L, 14' 0" D 620,863 600 1000 430 840 SS316 

Total 

Pressure Swing Adsorber Fixed Bed 12 25' 0" L, 10' 0" D 161,508 600 150 430 130 SS316 

$75,000,850 

Exchangers 

Gas Trim Condenser Shell/Tube 1 50 MMBTU/hr Duty 600 600 430 525 SS316 

Condensate Heater Shell/Tube 1 396.2 MMBTU/hr Duty 600 600 430 525 SS316 

SS316 (shell), 

Process Gas Cooler Shell/Tube 1 659.9 MMBTU/hr Duty 600 1000 430 1000 Haynes 556 (tubes) 

BFW Preheater Shell/Tube 1 306.5 MMBTU/hr Duty 600 600 430 525 SS316 

Total $23,214,575 

Pumps, Compressors & Blowers 

Natural Gas Compressor 

Total 

Water Pump Pos. Displacement 3 1839 GPM 600 100 430 70 683 HP CS 

$570,000 

SUBTOTAL $98,785,425 

TOTAL PROJECT COST $244,000,000 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Depictions of typical large scale SMR reactors and how they are integrated with fired heaters 

have been provided by Udhe and can be seen in Figures 2-4 and 2-512. These designs are more 

unique than what is used in small scale designs. While this requires additional capital for plant 

construction, the net result is a more efficient facility. 

 

FIGURE 2-4 TYPICAL LARGE SCALE 

SMR REACTOR DESIGN 

FIGURE 2-5 LARGE SCALE 

SMR REACTOR HEAT INTEGRATION 

A small scale SMR will typically purchase natural gas from a utility at a pressure slightly above 

atmospheric pressure. The large systems will purchase their natural gas from a pipeline 

transmission company or directly from the producer, using the transmission company as a 

common carrier. In these systems, the pipeline pressures are 200 to 1500 psig. Therefore, a 

natural gas compressor is not included in this design. While a water pump is included in the 

large scale equipment list, a similar situation may occur where industrial water is available at 

pressure, making a pump unnecessary. 

Table 2-6 highlights the main operating conditions for both the large and small scale cases. In 

general, the economy of scale in the large scale case allows more severe conditions, increasing 

the overall methane conversion and hydrogen recovery. 

 

12 Hydrogen, UHDE, ThyssenKrupp, http://www.uhde.biz/cgi-bin/byteserver.pl/pdf/broschueren/Duengemittel/Hydrogen.pdf (August 2005) 

Task 1 Final Report 2-13 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
TABLE 2-6 MAJOR PROCESS CONDITIONS, LARGE AND SMALL SMR DESIGNS 

Process Conditions Small Scale Design Large Scale Design 

Reforming pressure, atm 20 30 

Steam to carbon ratio 3 2.8 

Reformer inlet temp., °F 850 1000 

Reformer outlet temp., °F 1470 1615 

Methane conversion, % 71 77 

Shift inlet temp., °F 840 840 

Shift outlet temp., °F 480 525 

CO concentration in shifted gas, % 0.3 0.8 

PSA hydrogen recovery, % 75 86 

Natural gas feed/hydrogen product, vol. ratio 0.47 0.36 

Steam/hydrogen product, vol. ratio 1.4 0.97 

2.4 OPERATING COSTS AND UTILITY REQUIREMENTS 

The basis for the operating costs and utility requirements are shown in Table 2-7. 

TABLE 2-7 SMR OPERATING COSTS AND UTILITY REQUIREMENTS 

Operating Variable Large Scale Small Scale 

Cooling Water (GPM) 1839 N/A 

Natural Gas (MMSCFD) 15 0.002 

Catalyst 

 HDS/ZnO 

 Reforming 

 Shift 

 PSA Adsorbent

1,415 ft3/3,304 ft3 

7,300 ft3

9,983 ft3

250,000 ft3

0.06 ft3/0.14 ft3 

0.38 ft3

0.5 ft3

21 ft3

Electricity Use (kW) 9517 16 

The greater heat integration used in the large scale case leads to different utility requirements. 

For example, the small scale case uses mostly air-cooled exchangers, while the large scale case 

uses cooling water that is turned into steam for feed to the reformer reactor. Natural gas usage is 

predominantly for heat requirements in the reformer. Total electricity use comprises all power 

necessary for fans, blowers, motors, and starters. 

Similar catalyst are used in both cases, with Co/W HDS catalyst, Fe/CrO shift catalyst, and, 

typically, reforming catalyst that is Ru/Ni/AlO. Catalyst costs and lifespans will likely be the 

same in both cases. 

Task 1 Final Report 2-14 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Section 3 Natural Gas Liquid Expanders 

3.1 INTRODUCTION AND METHODOLOGY 

Detailed large and small scale Natural Gas Liquid (NGL) Expander designs were developed 

assuming a super-rich (<85% methane) NGL feed. Providing detail for a super-rich design 

demonstrates to the full extent the amount of equipment necessary to perform an extensive 

fractionation on a multi-component system. As the feedstock becomes leaner, both the size of 

the plant and the amount of equipment necessary for separation decreases. The small scale 

design assumes 75 MMSCFD of inlet gas, while the large scale design has 400 MMSCFD of 

inlet gas. These are typical unit sizes for the small and large ranges of the NGL expander design 

spectrum. 

The amount of ethane that is left in the methane stream depends largely on pipeline requirements 

and product economics. Current plants are designed to swing between ethane rejection (removed 

ethane is returned to the residual or product gas) and ethane recovery (removed ethane is 

included with the NGL product) depending on the operating conditions and how flows are 

directed in the plant. The small scale unit described below was based on ethane rejection, while 

the large scale unit recovers ethane. Regardless of the operating mode, the plant design will not 

change13. 

Data was collected from publications and licensors of NGL technology. The focus of this 

research was on the capital and operating costs of large and small scale designs based on turboexpander technology. Since this study includes both large and small scale designs, 

configurations have been developed to reflect the design which is considered most common and 

economic for the unit size. The major difference between the designs is in the heat integration, 

with the larger-scale system utilizing more cross exchange of the hot and cold streams to 

improve overall heat efficiency. The large scale unit has also been modeled in HYSYS to 

confirm vendor information. 

3.2 TECHNOLOGY BACKGROUND 

Natural gas that is extracted from terrestrial sources consists predominantly of methane, and, to a 

lesser extent, a quantity of natural gas liquids. These are typically ethane and propane, but also 

contain some butane and lesser amounts of higher carbon content, aliphatic hydrocarbons. Table 

3-1 below shows compositions for a range of potential feedstocks to a NGL plant. 

 

13 A number of literature sources confirm this approach, including “Liquid Levels and Density Case histories with DP cell problems”,

Hydrocarbon Online, March 2001, and “Deepwater Production Drives Design of New Gulf Gas Plant”, Oil & Gas Journal, March 16, 1998. 

Task 1 Final Report 3-1 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
TABLE 3-1 TYPICAL NGL FEEDSTOCK COMPOSITIONS 

Classification 

Super-Rich Rich Lean Super-Lean 

C1 (mole%) 82.00 87.00 92.00 96.30 

C2 9.70 7.90 4.80 2.20 

C3 4.50 3.30 1.80 0.80 

IC4 0.50 0.46 0.49 0.20 

NC4 1.80 1.13 0.87 0.50 

IC5 0.50 0.00 0.00 0.00 

NC5 0.49 0.21 0.04 0.00 

C6+ 0.51 0.00 0.00 0.00 

BTU/scf 1251.81 1161.77 1105.35 1054.51 

If the quantities of these other materials are low, they slightly enhance the heating value of the 

natural gas and may be left in the gas. If the amount of these materials is such that there may be 

condensation in the transmission pipeline, or there could be an impact on the suitability of the 

gas for downstream applications (usually measured in terms of the heating value and/or the 

related Wobbe number14), the pipelines may not purchase the gas from producers without prior 

treatment to remove the heavy hydrocarbons. 

In addition, a large industry has grown to utilize the ethane/propane and heavier hydrocarbons 

being removed from the natural gas as petrochemical feedstocks. Extraction of NGLs from 

natural gas can be either profitable or an unavoidable cost of production of natural gas for 

pipeline transport, depending on the relative price of the hydrocarbons on a heating value and 

chemical feedstock basis. 

Early facilities used circulating heavy hydrocarbon liquids to absorb the lighter components in 

lean oil absorber plants. In the 1960s, advances in turbomachinery led to newer cryogenic 

processes being introduced. These turboexpander processes have become the dominant 

technology. 

Technology has evolved to enable the expander plants to have reasonable recoveries and 

economics when rejecting ethane (leaving it in the gas) while recovering 90% or more of the 

propane and heavier components. This is necessary to enable such plants to remain in operation 

when economics favor leaving the ethane in the pipeline gas. There are a number of patented 

schemes that improve upon the original turboexpander design in order to enable this mode of 

operation. Many companies offer NGL expander plants including Fluor, Ortloff, IPSI LLC, 

ProQuip, and ABB Randall. 

 

14 The Wobbe number, or Wobbe index, of a fuel gas is found by dividing the high heating value of the gas in BTU per standard cubic foot by 

the square root of its specific gravity with respect to air. The higher a gases' Wobbe number, the greater the heating value of the quantity of 

gas that will flow though a hole of a given size in a given amount of time. The usefulness of the Wobbe number is that for any given orifice, 

all gas mixtures that have the same Wobbe number will deliver the same amount of heat. 

Task 1 Final Report 3-2 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Task 1 Final Report 3-3 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 

3.3 PROCESS DESCRIPTIONS AND DESIGN ASSUMPTIONS 

As with the other sections, flow diagrams, stream properties, and an equipment list for both cases 

is presented below. The multipliers used for installation and indirect costs for the SMR units 

have also been applied here. After calculation of the equipment cost, a 2.47 factor was used to 

calculate the Total Installed Cost. For indirect costs, factors developed by NREL to account for 

engineering costs, construction, contingencies, and fees were applied. Indirect cost factors were 

then added to the TIC to obtain the Total Project Investment. A summary of the results for both 

cases can be seen in Table 3-2. 

TABLE 3-2 NGL EXPANDER PROJECT COSTS, 2005 BASIS 

Large Scale Plant Small Scale Plant 

Feed Rate (MMSCFD) 400 75 

Equipment Cost ($MM) 18.8 4.8 

Installed Cost ($MM) 46.4 12.0 

Total Plant Investment ($MM) 62.4 16.1 

Total Investment/SCFD Feed ($) 0.16 0.21 

The differences between the small and large scale cases are not as pronounced as what was seen 

in the FCC and SMR cases. This is largely due to the relatively small size differences here when 

compared to the other technologies. NGL expanders of a scale smaller than ~25 MMSCFD are 

not commercially available; typically, ~75 MMSCFD is considered the low end of what is 

considered for commercial design. Recovering natural gas liquids from smaller streams would 

be performed using other technologies, such as straight refrigeration units or Joule-Thomson 

plants. As can be seen from Table 3-2, there is not a major economy of scale advantage between 

the high and low end of the commercial plant range. 

3.3.1 Small Scale NGL Expander Design 

Figure 3-1 shows the PFD for the small scale NGL expander, while Table 3-3 shows the stream

conditions. The equipment list for the large scale design can be found in Table 3-4. 
201C

Gas/Gas

Exch

202C

Warm

 Gas/Liquid

Exch

204C

Cold Gas/Liq

Exch

2

201F

High Pressure 

Separator

Tower

201E

DEC2

3

204J

Expander

202F

Low Pressure

Separator

202J/JA Low Pressure Liq Pump

5

6

8

7

4

207C

Condenser II

206C

Condenser I

205J

Expander-Compressor

201J

Residue Compressor

17 18

V-6

11

16

203C

Chiller

12

203F

Reflux

Drum

203J/JA

Reflux Pump

9

10

15 13

208C

Comp Discharge

Cooler

101F

Inlet Separator

101DA 101DB

Mole Sieve Drying

103F

Dust Filter

1

401J/JA

Hot Oil

Pump

Hot Oil System

101C

Regen Gas 

Cooler

401F

Hot Oil 

Vessel

102F

Regen

Gas

Separator

FC

19

To Waste 

Treatment

302E

Refrig Suction Vessel

301J

RefrigComp

301C

Refrig

Conden

301F

Refrigerant Surge

Vessel

303F

Economizer

14

401B

Waste 

Heat

Unit

205C

Column

Reboiler

Turbine 

Exhaust

To Atm

P-116

Natural Gas

Residue Product Gas

FIGURE 3-1 SMALL-SCALE NGL EXPANDER PROCESS FLOW DIAGRAM 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
TABLE 3-3 SMALL SCALE EXPANDER STREAM PROPERTIES 

Stream No. 1 2 3 4 5 6 7 

Stream ID Feed Gas 

Dry Gas to 

NGL Unit 

To High 

Pressure 

Separator 

Liq from 

High 

Pressure 

Sep 

Vapor 

from High 

Pressure 

Sep 

Expander 

Outlet 

Liquid 

from Low 

Pressure 

Sep 

Methane (lb-mole/h) 6670.1 6670.1 6670.1 756.3 5913.7 5913.7 191.2 Ethane 741.1 741.1 741.1 250.4 490.7 490.7 148.9 Propane 436.4 436.4 436.4 235.6 183.9 182.9 134.9 i-Butane 82.3 82.3 82.3 60.5 21.9 21.9 19.9 n-Butane 148.2 148.2 148.2 116.7 31.6 31.6 29.8 i-Pentane 41.2 41.2 41.2 36.2 5.0 5.0 4.9 n-Pentane 65.9 65.9 65.9 59.5 6.4 6.4 6.3 n-Hexane 35.4 35.4 35.4 33.9 1.5 1.5 1.5 n-Heptane 10.7 10.7 10.7 10.5 0.2 0.2 0.2 n-Octane 3.3 3.3 3.3 3.3 0.0 0.0 0.0 H2O 10.1 0.0 0.0 0.0 0.0 0.0 0.0 Total lb-mole/h 8,244.7 8,234.6 8,234.6 1,562.9 6,654.9 6,653.9 537.6 Total lb/h 174,347.5 174,164.8 174,164.8 52,394.3 121,170.5 121,770.5 17,344.3 Std Gas Flow MMscf/d 75.0 74.9 74.9 14.2 60.5 60.5 4.9 Total BBL/d 34,273.1 34,260.6 34,260.6 8,156.9 26,103.7 26,103.7 2,815.6 Vapor Fraction 1.0 1.0 0.8 0.0 1.0 0.9 0.0 Pressure, psia 1000 980 965 965 965 340 340 Temperature, F 102.5 102.5 -10.0 -10.0 -10.0 -87.6 -87.6 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Stream No. 10 11 12 13 14 15 16 17 

Stream ID 

Expander 

Vapor 

from 

Condenser 

Liquid 

after 

Valve 

High 

Press 

Liq to 

Tower 

Vapor 

from 

Tower 

Tower 

Bottom 

Residue 

gas to HX 

Residue 

Gas to 

Exp/Comp 

Residue 

Compr 

Suction 

Methane (lb-mole/h) 5722.6 756.3 756.3 947.4 0.1 6670.0 6670.0 6670.0 Ethane 341.9 250.4 250.4 373.9 25.3 715.8 715.8 715.8 Propane 48.0 253.6 253.6 17.4 371.0 65.4 65.4 65.4 i-Butane 2.0 60.5 60.5 0.1 80.3 2.0 2.0 2.0 n-Butane 1.7 116.7 116.7 0.0 146.5 1.7 1.7 1.7 i-Pentane 0.1 36.2 36.2 0.0 41.1 0.1 0.1 0.1 n-Pentane 0.1 59.5 59.5 0.0 65.8 0.1 0.1 0.1 n-Hexane 0.0 33.9 33.9 0.0 35.4 0.0 0.0 0.0 n-Heptane 0.0 10.5 10.5 0.0 10.7 0.0 0.0 0.0 n-Octane 0.0 3.3 3.3 0.0 3.3 0.0 0.0 0.0 H2O 0.0 0.0 0.0 0.0 0.0 0.0 0.0 0.0 Total lb-mole/h 6,116.4 1,580.9 1,580.9 1,338.8 779.5 7,455.1 7,455.1 7,455.1 Total lb/h 104,426.2 52,394.3 52,394.3 27,219.8 42,518.7 131,646.0 131,646.0 131,646.0 Std Gas Flow MMscf/d 55.6 14.4 14.4 12.2 7.1 67.8 67.8 67.8 Total BBL/d 23,288.1 8,156.9 8,156.9 5,745.6 5,226.9 29,033.7 29,033.7 29,033.7 Vapor Fraction 1.0 0.4 0.7 1.0 0.0 1.0 1.0 1.0 Pressure, psia 335 370 365 340 350 335 325 435 Temperature, F -36.3 -46.0 90.0 -43.0 199.1 -38.1 90.0 137.9 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
TABLE 3-4 SMALL-SCALE NGL EXPANDER EQUIPMENT LIST 

Design Operating 

Quantity Price, total Total Installed 

Item No Description Type Per Train Size Weight P T P T Power Useage Materials (unistalled) Cost 

2002 Cost Index 

lbs PSIG oF PSIG oF (No.) HP $US (US $) 

Reactors 

101D Mole Sieve Dryer Packed Bed 2 66" D x 26' T/T 52,000 1,100 600 1,000 100 CS $51,375 

Total $51,375 Exchangers 

201C Gas/Gas Exchanger Shell/Tube UA = 703000 Aluminum $95,000 202C 

1 1,100 -20 1,000 -10 

Warm Gas/Liquid Exchanger Shell/Tube UA = 132000 Aluminum $114,000 

203C 

1 1,100 -20 1,000 -10 

Chiller Shell/Tube UA = 106000 1,100 Aluminum $85,500 

204C 

1 -20 1,000 -10 

Cold Gas/Liquid Exchanger Shell/Tube UA = 189000 1,100 Aluminum $120,000 

205C 

1 -20 1,000 -10 

Reboiler Shell/Tube 200 sq. ft. CS included in 201E 

206C 

1 470 -50 350 200 

Condenser I Shell/Tube UA = 21000 Aluminum $78,000 

207C 

1 470 -100 350 50 

Condenser II Shell/Tube UA = 94000 470 Aluminum $76,000 

208C 

1 -100 340 -40 

Comp. Discharge Cooler Air Fin 1 4290 sq ft 1,100 300 1,000 290 43 CS $100,000 

209C Product Cooler Air Fin 1 1270 sq ft 470 250 350 200 12 CS $30,000 

Total $698,500 Columns, Vessels & Tanks 

101F Inlet Filter Separator Vertical 1 1100 150 1000 100 CS $27,000 

103F Dust Filter Vertical 1 1100 200 1000 100 CS $20,000 201F High Pressure Separator Horizontal 1 48" x 10' T/T 12,000 1100 -20 1000 -10 CS $148,540 202F Low Pressure Separator Horizontal 1 42" x 8' T/T 3,000 470 -100 350 -85 SS $98,522 

203F Reflux Drum Horizontal 1 48" x 10' T/T 4,700 470 -50 340 -40 CS included in 201E 

201E Deethanizer Vertical 1 52" x 40' T/T 26,000 470 -50 350 50 LTCS $682,073 

Total $976,135 Pumps, Compressors & Blowers 

Expander 

202J 204J Turbine 1 1217 HP 400 -100 350 -85 1,217 CS/SS $565,685 

Low Pressure Liquid Pump 

203J 

Centrifugal 2 70 gpm x 50 psig 470 -100 350 -85 3 SS $108,902 

Reflux Pump 

205J 

Centrifugal 2 95 gpm x 50 psig 470 -50 340 -40 5 CS $48,099 

Expander Compressor 

201J 

Turbo-Expander 1 1217 HP 1,100 150 435 130 1,217 CS included with 204J 

Residue Compressor Centrifugal 1 4113 HP 1,100 150 1,000 290 4,113 CS $2,375,879 

206J Product Pump Centrifugal 2 98 HP 470 250 350 200 98 CS $40,000 

Total $3,138,565 Other 

Refrigeration System 180 Included in 203C 

Hot Oil System 21 Included in 201E 

SUBTOTAL $4,864,575 

$12,015,501 TOTAL PROJECT COST 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
The feed gas, stream 1, at 1000 psia enters the inlet filter separator (101F) where any entrained 

liquids and solids are removed. The gas, free of entrained liquids, flows to the mole sieve dryer 

(101DA or 101DB) where the water is adsorbed on the mole sieves, reducing the water content 

of the gas to less than 1 ppm. The mole sieve beds are regenerated by bringing turbine exhaust 

gas (which is predominantly methane) from the residue compressor (201J) through the waste 

heat unit (401B) where a portion of the gas from the NGL recovery section is heated to 550°F 

prior to flowing through the switching valves into the mole sieve dryer. The gas then enters the 

dust filter (103F) where any mole sieve fines are removed. From the dust filter, the dry gas, 

stream 2, enters the NGL recovery section of the plant. 

The dehydrated gas is split to act as a heat exchange medium for a number of different streams. 

46% of the gas goes through the gas/gas exchanger (201C), where it is cooled to -25°F against 

the residue gas. The remainder of the gas flows through the warm gas/liquid exchanger (202C), 

the chiller (203C), and the cold gas/liquid exchanger (204C) to cool the stream to 4°F. These two 

split streams are recombined, and the combined stream, stream 3, enters the high pressure 

separator (201F) where the cold liquids are separated from the gas. The cold liquids, stream 4, 

leave the separator and are reduced in pressure across a valve to reduce the temperature to -45°F. 

Gas stream 5, from the high pressure separator (201F) is expanded almost isentropically in the 

expander (204J) to 340 psia. This gas stream, stream 6, enters the low pressure separator (202F) 

where the liquids are separated from the gas. These cold liquids, stream 7, are pumped by the 

low pressure liquid pump (202J, JA) through condenser I (206C) to the deethanizer (201E). The 

gas from the low pressure separator, stream 8, flows through condenser II (207C) before being 

recombined with the deethanizer overhead, stream 10. 

The deethanizer (201E) is a packed tower which separates the propane and heavier components 

from the methane and ethane. Heat is supplied to the deethanizer by a hot-oil-heated reboiler 

(205C). The overhead vapors from the deethanizer are partially condensed in condenser I (206C) 

and condenser II (207C) using the cold gases and liquids from the low pressure separator (202F). 

The liquids are separated in the reflux drum (203F) and pumped back to the deethanizer by the 

reflux pump (203J,JA). 

Stream 13 is combined with stream 10 from the low pressure separator (202F) and used for 

cooling in the gas/gas exchanger (201C). This gas, stream 16, is partially re-compressed by the 

expander compressor (205J), which uses the power generated by the expander (204J). The gas, 

stream 17, is further compressed to pipeline pressure by the gas turbine driven residue 

compressor (201 J). The residue compressor discharge, stream 18, is cooled by the compressor 

discharge cooler (208C) prior to entering the pipeline. 

External refrigeration is provided to the chiller (203C) by a propane refrigeration system. The 

gas engine driven refrigerant compressor (301J) compresses the propane to 245 psia after which 

it is condensed using air in the refrigerant condenser (301C). The liquid propane then goes to the 

refrigerant surge vessel (301F) and on to the economizer (303F) where it is flashed. The 

economizer serves to reduce the horsepower requirements of the refrigeration system. The 

vapors from the economizer are returned interstage into the refrigerant compressor. The liquid 

from the economizer goes to the chiller (203C) where it vaporizes to provide cooling for the inlet 

gas to the NGL recovery section of the expander plant. The vapors from the chiller go through 

Task 1 Final Report 3-8 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
the refrigerant suction vessel (302F) to remove any entrained liquid and are then returned to the 

suction of the refrigerant compressor. 

A hot oil system provides energy for the regeneration of the dryers. Heat for the mole sieve 

regeneration gas and the hot oil is provided by the turbine exhaust from the residue gas 

compressor (201J). This heat is recovered in the waste heat unit (401B). Hot oil is pumped 

through this unit by the hot oil pump (401J, JA) to the deethanizer column reboiler (205C). Hot 

oil is returned from the reboiler to the suction of the hot oil pump. 

Design Assumptions 

The small scale NGL design was based on a study performed by M. W. Kellogg for GRI15. A 

self-contained, turbo-expander plant for the recovery of NGL from natural gas was designed and 

costed. Standard sizing techniques were used to estimate the size and weight of process 

equipment. Telephone estimates were received from vendors for most equipment. Bulks for the 

process units were determined using a Kellogg-proprietary estimating program with adjustments 

for the use of skid-mounted units. The costs were updated to 2005 using the Chemical 

Engineering Plant Cost Index. 

The design presented here assumes a stand-alone facility that could be used for most NGL 

expander designs. The feed natural gas entering the plant is assumed to require additional 

treating, which is why both drying and particulate removal units have been included. While 

these steps may not always be necessary for a unit placed on a common natural gas header, 

including them gives the plant extra flexibility to handle a variety of feeds. 

Heating and cooling loads in an expander facility may vary considerably based on the overall 

feed rate, pipeline quality requirements, and the level of fractionation desired. For this design, 

both hot oil and refrigeration systems have been included to provide additional plant flexibility. 

These units are typically seen more often in small units rather than large ones which have more 

stable plant operations. 

3.3.2 Large Scale NGL Expander Design 

Figure 3-2 shows the PFD for the small scale NGL expander, while Table 3-5 shows the stream 

conditions. The equipment list for the large scale design can be found in Table 3-6. 

The large scale process for NGL recovery is very similar to that of the smaller scale systems 

using the same turbo-expander technology. The major difference is heat integration—the larger 

scale systems are able to afford additional heat exchanger systems to decrease the 

thermodynamic penalties related to cooling the stream. This adds to the overall complexity of 

the design; while increasing capital costs, the total investment per SCFD of feed decreases. 

Inlet gas enters the plant at 120°F and 900 psia (stream 1). If the inlet gas contains a 

concentration of sulfur compounds which would prevent the product streams from meeting 

specifications, the sulfur compounds are removed by appropriate pretreatment of the feed gas. In 

addition, the feed stream is usually dehydrated upstream to prevent hydrate (ice) formation under 

 

15 Evaluation of Natural Gas Processing Technologies, Topical Report, Task 3, Volume II, M. K. Changela, et al, The M. W. Kellogg 

Company,(1990). 

Task 1 Final Report 3-9 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Task 1 Final Report 3-10 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 

cryogenic conditions. Solid desiccants are typically used for this purpose and are not shown 

here. 

The feed is cooled in several heat exchangers with cool residue gas at –86°F, bottom liquid 

product at 62°F, demethanizer reboiler liquids at 39°F, and demethanizer side reboiler liquids at -

32°F. The cooled stream 3 enters high pressure separator (V-100) at -22°F and 885 psia where 

the vapor (stream 5) is separated from the condensed liquid (stream 4). 

The vapor from separator V-100 (stream 5) enters the turbo-expander (K-100) in which 

mechanical energy is extracted from this portion of the high pressure feed. This unit expands the 

vapor isentropically from a pressure of 885 psia to a pressure of 312 psia, with the work 

expansion cooling the expanded stream 6 to a temperature of approximately -112°F. Typical 

commercially available expanders are capable of recovering on the order of 80-85% of the work 

theoretically available in an ideal isentropic expansion. The work recovered is used to drive the 

centrifugal expander compressor (K-101). The expanded and partially condensed stream 6 is 

further flashed in the low-pressure separator (V-101). 
Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 

FIGURE 3-2 LARGE-SCALE NGL EXPANDER PROCESS FLOW DIAGRAM 
Task 1 Final Report 3-12 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 

TABLE 3-5 LARGE SCALE EXPANDER STREAM PROPERTIES 

Name 1 2 3 4 5 

Methane [lbmole/h] 36014.15 36014.15 36014.15 6272.595 29755.73 

Ethane 4260.211 4260.211 4260.211 2042.24 2223.505 

Propane 1976.387 1976.387 1976.387 1436.301 523.1776 

n-Butane 1668.949 1668.949 1668.949 1481.185 196.1916 

Water 0 0 0 0 0 

Molar Flow [lbmole/h] 43919.7 43919.7 43919.7 11221.1 32698.6 

Vapor Fraction 1.00 1.00 0.82 0.00 1.00 

Temperature [F] 120.0 72.0 -22.0 -38.6 -38.6 

Pressure [psia] 900 900 900 885 885 

Mass Flow [lb/h] 890,033.3 890,033.3 890,033.3 311,080.5 578,952.9 

Std Ideal Liquid Flow [barrels/d] 179,967.1 179,967.1 179,967.1 53,441.8 126,525.3 

Std Gas Flow [MMSCFD] 400.0 400.0 400.0 102.2 297.8 

Name 6 7 8 9 11 

Methane [lbmole/h] 29755.73 1262.272 28487.46 1262.272 6272.595 

Ethane 2223.505 958.1792 1252.852 958.1792 2042.24 

Propane 523.1776 461.8768 89.4894 461.8768 1436.301 

n-Butane 196.1916 189.3408 0 189.3408 1481.185 

Water 0 0 0 0 0 

Molar Flow [lbmole/h] 32698.6 2868.8 29829.8 2868.8 11221.1 

Vapor Fraction 0.91 0.00 1.00 0.00 0.40 

Temperature [F] -112.0 -112.0 -112.0 -111.9 -83.5 

Pressure [psia] 312 312 312 324 325 

Mass Flow [lb/h] 578,952.9 80,332.8 498,620.1 80,332.8 311,080.5 

Std Ideal Liquid Flow [barrels/d] 126,525.3 14,207.6 112,317.6 14,207.6 53,441.8 

Std Gas Flow [MMSCFD] 297.8 26.1 271.7 26.1 102.2 

Name 14 15 16 17 18 19 

Methane [lbmole/h] 205.3818 45039.73 1345.113 1352.006 45045 35810.74 

Ethane 4044.442 271.8696 3301.641 3301.884 271.9014 216.1614 

Propane 1974.825 0 1202.45 1202.538 0 0 

n-Butane 1674.652 0 937.503 937.572 0 0 

Water 0 0 0 0 0 0 

Molar Flow [lbmole/h] 7899.3 45311.6 6793.5 6794.0 45316.9 36026.9 

Vapor Fraction 0.00 1.00 0.00 0.46 1.00 1.00 

Temperature [F] 115.0 -155.1 -47.5 38.0 288.5 90.0 

Pressure [psia] 600 300 314 314 900 900 

Mass Flow [lb/h] 309,317.9 730,573.2 228,587.0 228,570.1 730,655.8 580,871.3 

Std Ideal Liquid Flow [barrels/d] 47,364.8 166,802.7 37,646.1 37,644.6 166,821.9 132,623.4 

Std Gas Flow [MMSCFD] 71.9 412.7 61.9 61.9 412.7 328.1 
TABLE 3-6 LARGE-SCALE NGL EXPANDER EQUIPMENT LIST 

Design Operating 

Quantity Price, total Total Installed 

Item No Description Type Per Train Size Weight P T P T Power Useage Materials (unistalled) 

2002 Cost Index 

Cost 

lbs PSIG oF PSIG oF (No.) HP $US (US $) 

Exchangers E-100 Cold Gas/Gas Exchanger Shell/Tube 1 Area = 10,120 sq. ft. 350 -155 312 -112 Aluminum $179,197 

E-101 Condenser I Shell/Tube 1 Area = 16,145 sq. ft 1,000 -70 300 -60 Aluminum $174,602 

E-102 Condenser II Shell/Tube 1 Area = 15,725 sq. ft 1,000 -115 300 -110 Aluminum $222,847 

E-103 Condenser III Shell/Tube 1 Area = 17,990 sq. ft. 1,000 -150 300 -150 Aluminum $192,981 

E-105 Compressor Discharge Cooler Shell/Tube 1 Area = 17,480 sq. ft. 1,000 300 900 288 CS $229,740 

E-106 Gas/Liquid Exchanger Shell/Tube 1 Area = 24,025 sq. ft. 1,000 150 900 120 CS $261,903 E-107 Warm Gas/Gas Exchanger Shell/Tube 1 Area = 8,220 sq. ft. 1,000 150 900 120 CS $218,253 

E-108 Gas/Liquid Exchanger Shell/Tube 1 Area = 22,245 sq. ft. 1,000 100 900 72 CS $275,688 

E-109 Cold Gas/Gas Exchanger Shell/Tube 1 40 MMBTU/hr 1,000 -100 900 -20 Aluminum $252,714 E-110 Cold Gas/Liquid Exchanger Shell/Tube 1 Area = 13,324 sq. ft. 1,000 -50 350 -20 Aluminum $275,688 Total $2,283,613 Columns, Vessels & Tanks T-100 Demethanizer Tray Column 1 78" x 85' T/T 400 -160/60 300 -155/30 LTCS $1,033,829 

V-100 High Pressure Separator Horizontal 1 72" x 15' T/T 68,000 1000 -40 900 -22 CS $225,145 

V-101 Low Pressure Separator Horizontal 1 63" x 12' T/T 102,000 400 -120 312 -112 SS $149,331 

Total $1,408,305 Pumps, Compressors & Blowers P-100 NGL Pump Centrifugal 2 400 HP 650 -120 350 -100 400 SS $275,688 

P-101 Reflux Pump Centrifugal 2 12 HP 350 -120 350 115 12 SS $121,762 

K-100 Expander Turbine 1 5180 HP 1,000 -120 900 -85 5,180 CS $2,828,427 K-101 Expander Compressor Turbo-Expander 1 5180 HP 1,000 150 314 -50 5,180 CS included with K-100 

K-102 Residue Compressor Centrifugal 1 26,624 HP 1,000 150 900 288 26,624 CS $11,879,394 

Total $15,105,271 

SUBTOTAL $18,797,189 

TOTAL PROJECT COST $46,429,057 

Task 1 Final Report 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
The vapor (stream 8) from V-101 is divided into two streams. About 60% of the total vapor, 

passes through the cold gas/gas exchanger (E-100) against a portion of the cold distillation 

stream, resulting in further cooling and partial condensation. The further cooled and partially 

condensed stream is then supplied to the distillation column at an upper mid-column feed point. 

The remaining 40% of the vapor from V-101 is fed to the demethanation tower (T-100) at a 

second upper mid-column feed point. The condensed liquid (stream 9) from V-101 is pumped in 

the NGL pump (P-100) prior to feeding the distillation column at a third upper mid-column feed 

point. 

A portion of the high-pressure residue gas is withdrawn from the main residue flow (stream 18) 

to become the top distillation column feed. The recycle gas stream passes through a series of heat 

exchangers with a portion of the cool residue gas where it is cooled to -148°F. As the stream is 

expanded to the demethanizer’s operating pressure of 324 psia, it is cooled to a temperature of 

approximately -158°F. 

The cooling and partial condensation of a portion of the vapor phase leaving the work expansion 

machine results in additional liquids being supplied to the fractionation tower at the upper midcolumn feed position, whereupon these liquids act as reflux on the vapor rising up the tower. 

This in turn reduces the amount of vapor to be rectified by the top reflux stream (and also creates 

a leaner vapor composition), allowing a corresponding reduction in the amount of top reflux feed 

required to achieve the desired C2+ component recovery level. 

Design Assumptions 

The large scale NGL design was based on a study performed by GTI for DOE. A turboexpander plant for the recovery of NGL from natural gas was designed and costed. A HYSYS 

simulation model was used to size equipment. Costs were verified by Trimeric Corporation 

using vendor quotes and PDQ$ costing program. 

Large scale units have stable feed rates and qualities, since they are typically built in an area with 

a very steady feed supply. In addition, product demand is usually locked in and integrated with 

petrochemical facilities. The design flexibility incorporated into the small scale unit, processes 

for feed treating and additional heating/cooling, is typically not necessary in a large scale unit. 

3.4 OPERATING COSTS AND UTILITY REQUIREMENTS 

The basis for the operating costs and utility requirements are shown in Table 3-7. 

TABLE 3-7 NGL EXPANDER OPERATING COSTS AND UTILITY REQUIREMENTS 

Operating Variable Large Scale Small Scale Notes 

Cooling Water (gal/hr) 52,568 10,000 Interstage compressor requirements 

Catalyst 

 Mol Sieve (lb) N/A 19,500 ~$2.50/lb, lifespan will vary 

Electricity Use (MW) 27.9 5.1 Product compression is main cost 

Task 1 Final Report 3-14 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Besides the items listed above, natural gas is also used for various heating requirements, such as 

reboilers and hot oil system reheating. Typically, the fuel gas produced from the expander 

system is used for this service, so no external natural gas is purchased specifically for plant use. 

As mentioned earlier, the feed natural gas may require additional treating, which is why both 

drying and particulate clean-up units have been included in the small scale design. Large scale 

units which have more stable and reliable feed sources may have treating upstream of the facility 

and may not require these units to be present. Information provided by vendors for the large 

scale designs did not include these minor pieces of equipment, which is why no molecular sieve 

requirement is listed for the large scale unit. 

Task 1 Final Report 3-15 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Section 4 Summary and Conclusions 

When attempting to draw conclusions based on the scaling factors for plant sizes and costs 

addressed here, it should be recognized that the plant ranges considered are very different. This 

is important not only because of the sheer magnitude difference, but also for where the plants are 

situated in the range of commercial sizes. For example, comparing plants at the very low and 

high end of their respective commercial ranges will differ from a comparison between plants at 

other points along the commercial spectrum. This is compounded by the fact that the 

“commercial range” is a moving target based on technology advances and the industry where the 

technology is used. 

Most technologies have some asymptotic economy of scale when capital investment is compared 

to plant size. This is demonstrated in Figure 4-1 below, an evaluation of the economy of scale 

advantage for large scale SMR technology. 

FIGURE 4-1 SMR PROJECT COSTS, BROAD SIZE RANGE16

As can be seen from Figure 4-1, the economy of scale diminishes as the plant size approaches the 

high end of the commercial scale. When comparing two different plant sizes, it is important to 

consider at what points along this curve the plants are situated. While a power relationship is 

helpful in performing sizing analysis regardless of where the two plants are in relation to one 

another, more accuracy can be gained if the power is modified based on the relative position to 

each other on the curve. 

For the technologies evaluated in this study, different commercial size ranges were evaluated. 

The ranges are: 

 

16 The Hydrogen Economy: Opportunities, Costs, Barriers, and R&D Needs (2004), National Academy of Engineering, Board on Energy and 

Environmental Systems 

Task 1 Final Report 4-1 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
 FCC: Very low end commercial to mid-to-high end commercial. No commercial 

scale FCC units are built today below ~10,000 BPSD, due to the poor economics of 

units this size. 

 SMR: Very low end niche commercial unit to high end commercial unit for 

industrial purposes. The small scale unit would be used for hydrogen production for 

a small fuel cell application, while the large scale unit would be appropriate for a 

large industrial application. 

 NG Expander: Low-to-mid commercial to high end commercial. The range of 

commercial applications for NG expanders is lower than other technologies, due to 

technology breakpoints at the low end and a switch to multiple trains on the high end. 

Plant scale evaluations were limited to single train systems. 

Although the size ranges that were considered are different, an attempt was made to determine if 

a simple power relationship could be used to predict the cost of each small scale plant. The 

following basis was used: 

7.0

Capacity Scale Large

Scale Small Capacity Cost Scale Small CostScale Large ⎟

⎟

⎠

⎞

⎜

⎜

⎝

⎛ = ∗

The results of this analysis showed that at a minimum, an order of magnitude predictive basis is 

possible using this relationship for all technologies evaluated, regardless of the differences in 

size. In most cases, using a power basis provided a cost estimate within a level of acceptability 

(+/- 30%) for this level of detail. The analysis performed is detailed in Table 4-1 below. 

TABLE 4-1 DESIGN RATIOS AND PREDICTED COSTS, ALL TECHNOLOGIES 

FCC SMR 

NG 

Expander 

Large/Small Scale Size Ratio (Feedstock) 25 7,000 5.3 

Large/Small Scale Cost Ratio 9.4 1,000 3.8 

Size Ratio/Cost Ratio 2.7 7 1.4 

Predicted Small Cost, Power Basis ($MM) 20.2 0.75 19.3 

Actual Small Cost, Design Basis ($MM) 20.5 0.39 16.1 

Percent Difference 1.5% 92.3% 19.8% 

Although applying a power relationship to SMR reforming technology had the worst predictive 

capacity, the resulting estimate is still very close on an absolute basis. It was expected that this 

technology would have the most difficulty in being able to predict by a power relationship 

because of the considerable difference in size relative to the other technologies chosen. 

This power relationship can be used for many different types of technologies to provide an 

acceptable cost estimate for planning purposes. However, two major items should be kept in 

mind as this approximation is applied. First, the level of technological maturity will have an 

Task 1 Final Report 4-2 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
impact on how the costs scale when the plant size is changed. All operations being analyzed in 

this study are commercially mature, with little future cost reduction expected from technological 

breakthroughs. The impacts of future cost reductions due to technical improvements must be 

considered separately outside of this analysis. Secondly, the power that should be used when 

comparing two units will vary based upon their relative size. Comparing large scale units in the 

“flat” portion of Table 4-1 will be better accomplished with a larger exponent (reflecting a more 

linear relationship), while a smaller exponent should be used when comparing units that have a 

more asymptotic relationship. 

A few general conclusions can be made regarding the cost relationship as technologies are 

scaled: 

 The power relationship that should be used depends largely on the size range that is 

being considered. For example, a very large difference in sizes (SMR technology) 

will usually be much more difficult to accurately predict than smaller differences. 

 Economies of scale are usually greater when technologies contain a novel or 

expensive main piece of equipment, such as the catalytic furnace in an SMR unit or a 

reactor/regenerator unit in an overall FCC design. Established technologies, such as 

refrigeration and fractionation in the NG Expander designs, are more easily predicted 

as to their impact when sizes change. 

 Many technologies have a breakpoint where the type of design changes due to size 

factors. This is true for all technologies considered in this study, as well as for many 

other technologies such as oxygen production units. As designs get close to this 

technology breakpoint, the capital cost per unit of feed changes considerably. 

The power relationships explored here could be applied to technologies of particular interest to 

NREL, such as biomass gasifiers and tar crackers. Since neither of these technologies is mature, 

estimates for large scale applications will need to take into account the cost reductions expected 

as the technology progresses toward commercialization. Scaling from the small scale units 

currently operated or being explored by NREL to large, commercial units would likely require a 

smaller power relationship than the one used in Table 4-1. Any attempt to estimate unit costs 

from scaling factors should be checked with a preliminary (+/- 30%) design and cost estimate, 

using design software and vendor quotes. This step is important to provide a reality check for 

units that do not have widespread commercial application, and to provide additional data to 

support theoretical predictions. 

Task 1 Final Report 4-3 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Appendix A Technology Evaluations 

After discussions with NREL, it was decided that Subtask 1.1, Technology Evaluation, would 

focus on Chemical Process Technologies. The final list of technologies chosen was: 

 Refinery Processes 

 Claus Sulfur Recovery 

 Natural Gas Reforming to Hydrogen 

 Methanol from Natural Gas 

 Ammonia Production 

 Gas-To-Liquids 

 Natural Gas Liquids Expanders 

 Cryogenic Oxygen Plants 

 Ethylene and Olefins Plants 

Information about all of these technologies, including history, technology types, a description of 

the main technologies used, and commercial process ranges, was produced as the deliverable for 

Subtask 1.1. The main content of that deliverable is presented in this Appendix. 

After a review of Subtask 1.1, three technologies, Fluidized Catalytic Cracking of refinery gas 

oil, Natural Gas Reforming, and Natural Gas Liquids Expanders, were chosen for further 

investigation. These technologies were selected due to how the relationship between cost and 

equipment size at both a large and small scale is useful for application toward other technologies 

of interest to NREL. 

A.1 REFINERY PROCESSES 

A.1.1 Introduction 

As with many commodity chemical processes, petroleum refineries have increased in scale 

considerably since the first ones were built in the early 20th century. Economies of scale have 

played a large part in defining the current make-up of petroleum refineries in the US today. Due 

in part to the expansion of distribution facilities, environmental regulations, and the removal of 

price controls in 1981, many of the very small scale (less than 10,000 barrels per day (BPD)) 

refineries have shutdown since 1980. The total number of refineries has decreased from over 

300 in 1980 to 144 in 200417,18. 

 

17 2004 data from the US Department of Energy EIA-820 Refinery Capacity Survey, January 1, 2004. 

18 Historic information from the DOE commissioned report by the National Petroleum Council, “US Petroleum Refining, Assuring the 

Adequacy and Affordability of Cleaner Fuels”, June 2000. 

Task 1 Final Report A-1 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
A.1.2 Overall Refinery Sizes 

While considerable range still exists in US petroleum refineries, many very small scale refineries 

are filling either a geographic or specialty product niche. Of the 144 refineries, only 11 (8%) 

have a capacity under 10,000 BPD, while 62 (43%) have capacities over 100,000 BPD. Figure 

A-1 below has a breakdown of current refinery capacities: 

400-499 kBD

2%

300-399 kBD

4%

200-299 kBD

15%

100-199 kBD

21%

50-99 kBD

28%

<49 kBD

29%

500 kBD+

1%

FIGURE A-1 2004 US REFINERY SIZES 

While refineries less than 1,000 BPD have historically been in commercial operation, no 

refineries of this size are currently in operation in the US. The difference between the largest 

and smallest 2004 nameplate capacities is a factor of nearly 300. Table A-1 below shows the 

largest and smallest US refinery capacities as of January 1, 200419. 

Table A-1 Largest/Smallest US Refineries 

 (Nameplate Calendar Day Capacity, January 2004) 

Company Location Capacity (BPD) Major Outputs 

LARGEST 

ExxonMobil Baytown, TX 557,000 Fuels, Chemicals, Lubricants 

Hovensa LLC Kings Hill, USVI 495,000 Fuels, Chemicals, Lubricants 

ExxonMobil Baton Rouge, LA 493,500 Fuels, Chemicals, Lubricants 

SMALLEST 

Foreland Refining Eagle Springs, NV 1,925 Asphalt, Bunker Fuels 

Tenby Inc. Oxnard, CA 2,800 Asphalt, Bunker Fuels 

Silver Eagle Refining Evanston, WY 3,000 Asphalt, On-Road Fuels 

With every refinery having a different process flow and complexity level based on the crude 

types fed and the product slate chosen, it is difficult to make general comparisons between all 

refineries. A 500,000 BPD refinery making a variety of products will look very different from a 

2,000 BPD refinery that produces asphalt and bunker fuels. Large refineries will tend to have a 

greater level of complexity to optimize process economics by upgrading low quality crudes into 

 

19 All refinery process data from the US Department of Energy EIA-820 Refinery Capacity Survey, January 1, 2004. 

Task 1 Final Report A-2 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
valuable clean fuels. Smaller refineries cannot afford this level of complexity, focusing on a 

very limited input and output slate. 

A.1.3 Process Unit Sizes 

In order to provide comparisons on a more equal footing, petroleum refineries have been broken 

down by process unit type. Each of these process units are complicated processes by themselves, 

with many having been scaled considerably. Process size ranges are provided for each unit type. 

Please note that the capacities are for individual units at the plant and not the total amount of 

capacity available at the site. For example, although ExxonMobil’s Baytown refinery has 

557,000 BPD of atmospheric distillation capacity, this is performed over a number of individual 

units. 

Atmospheric and Vacuum Crude Distillation 

The first step of any petroleum refinery is to feed crude oil into a distillation column to obtain the 

rough product cuts that will be further refined and blended downstream. Most initial distillation 

is done at atmospheric conditions. When feeding a heavy crude slate, bottoms from atmospheric 

distillation units are sometimes sent to a vacuum crude tower for further component separation. 

Table A-2 Atmospheric/Vacuum Crude Distillation Unit Capacities 

Company Location Capacity (BPD) 

LARGE 

ExxonMobil Baytown, TX 90,000 to 120,000 (atm.) 

ChevronTexaco Pascagoula, MS 70,000 to 100,000 (vac.) 

SMALL 

Foreland Refining Eagle Springs, NV 1,925 (atm) 

Wyoming Refining Newcastle, WY 1,800 (atm) 

Alkylation 

This process consists of the reaction of isobutane with a mixed light olefinic (usually C3 and C4) 

stream to produce a high octane gasoline blending component. The resulting product is usually 

blended to make premium, 90 to 93 octane, gasoline. This reaction occurs at cold temperatures 

and low pressures, using stirred sulfuric acid as a reaction catalyst. 

Table A-3 Alkylation Unit Capacities 

Company Location Capacity (BPD) 

LARGE 

BP Texas City, TX 20,000 to 30,000 

SMALL 

Montana Refining Great Falls, MT 800 

Aromatics 

In general, aromatics units tend to be pair with more complex refineries that have both reforming 

capacity and a strong market for aromatics products (benzene, toluene, and xylenes). Large 

Task 1 Final Report A-3 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
refineries that are paired with olefins plants also usually possess some sort of aromatics 

processing capacity. Raw feed from refinery reformers or heavy sections of olefins plants are 

sent to aromatics processing units for extraction. This is usually a physical conversion, which 

consists of solvents, zeolite adsorption, and distillation. 

Table A-4 Aromatics Unit Capacities 

Company Location Capacity (BPD) 

LARGE 

BP Texas City, TX 45,000 

SMALL 

Alon USA LP Big Spring, TX 1,000 

Fluidized Catalytic Cracking 

A standard process in many refineries is the upgrading of gas oil to gasoline. FCC units have 

been present in US refineries for over 50 years, and are considered a very mature technology. In 

this process, a fluidized catalyst reacts with an inlet gas oil stream at high pressure to produce a 

predominantly unsaturated product stream suitable for gasoline blending. The catalyst is 

separated from exit gases in a cyclone, regenerated in a separate reactor, and then reintroduced 

into the process reactor. 

Table A-5 Fluidized Catalytic Cracking Unit Capacities (Fresh Feed) 

Company Location Capacity (BPD) 

LARGE 

ExxonMobil Baytown, TX and Baton Rouge, LA 70,000 to 110,000 

SMALL 

Montana Refining Great Falls, MT 2,400 

Hydrocracking 

In this process, gas oil or distillate is converted to lighter, higher octane blending components in 

the presence of hydrogen. Unlike an FCC unit, the process occurs over a fixed bed at high 

pressure. Because of the presence of hydrogen in the reactor, the product produced is saturated, 

with different blending properties than FCC product. 

Table A-6 Hydrocracking Unit Capacities 

Company Location Capacity (BPD) 

LARGE 

ExxonMobil Beaumont, TX 25,000 to 35,000 

SMALL 

Wynnewood Refining Wynnewood, OK 5,500 

Task 1 Final Report A-4 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Naphtha Reforming 

Many straight-run pipestill naphthas or condensates from natural gas liquid processing have low 

octane values due to the presence of paraffinic hydrocarbons. In order to increase the octane 

value and make the naphtha streams more suitable for blending, reformers are used. Reforming 

reactions usually occur at high temperatures over fixed-bed platinum catalysts. The product 

reformate is a branched, unsaturated hydrocarbon stream. Hydrogen is also produced in this 

reaction. 

Table A-7 Naphtha Reforming Unit Capacities 

Company Location Capacity (BPD) 

LARGE 

BP Whiting, IN 30,000 to 45,000 

SMALL 

Somerset Refining Somerset, KY 1,000 

Desulfurization 

Unless the crude slate is very sweet, most refinery gasoline and on-road distillate products 

require desulfurization to meet product specifications. This is a mature technology, using 

hydrogen and a fixed-bed catalyst to remove sulfur from the product stream. 

Table A-8 Desulfurization Unit Capacities 

Company Location Capacity (BPD) 

LARGE 

ExxonMobil Baytown and Beaumont 30,000 to 40,000 

SMALL 

Montana Refining Great Falls, MT 1,000 

A.2 CLAUS SULFUR RECOVERY 

A.2.1 Introduction 

The Claus process is the principal process used for large scale (above ~25 TPD) sulfur recovery. 

The size of Claus plant used depends on the required degree of sulfur recovery, often dictated by 

air pollution regulations. The largest Claus plants are approaching 3000 TPD. 

For smaller capacities (~1 TPD or less), a scavenging bed containing minerals, such as zinc 

oxide or liquid hexahydrotriazines, that react with sulfur is the predominant method of removal. 

At the intermediate scale (~1 to 20 TPD), regenerable sulfur recovery processes such as liquid 

reduction-oxidation (Redox), are more often employed. Examples of these processes include 

Stretford, LO-CAT II, and SulFerox, where chelated metal catalysts are used to convert H2S 

directly to elemental sulfur. Economics make these plants attractive predominantly when the 

amount of sulfur removal is relatively low (under 20 TPD). 

The need to process increasingly sour crude oil in refineries and increasing amounts of subquality gas containing H2S will increase the need for future Claus plants. Technology 

improvements have mainly focused on tail gas treating processes to deal with ever tightening 

Task 1 Final Report A-5 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
sulfur recovery requirements and means of increasing throughput in existing Claus plants. The 

basic modified Claus process is essentially unchanged since its widespread introduction in the 

1950s. 

A.2.2 Process Description 

Claus plants are the major means of recovering sulfur from natural gas and refinery gas when 

processing requirements are ~25 TPD and above. The process involves removing the acid gas 

(CO2 and H2S) with an amine process or similar system upstream of the Claus unit, and then 

feeding the resulting acid gas from the amine stripper into a furnace (see Figure A-2). In the 

furnace, about 1/3 of the H2S is converted to SO2. The H2S and SO2 react to form elemental 

sulfur, with approximately 70% of the sulfur recovered in a condenser following the furnace. 

Multiple stages of catalyst beds with intermediate reheat are provided downstream of the Claus 

furnace to recover additional sulfur. Each stage reacts 2 moles of H2S with one mole of SO2

over a catalyst to produce elemental sulfur. A practical limit of 3 such stages has been found, 

since progressively less sulfur is removed in each stage. About 97% of the feed sulfur can be 

removed in this manner. If additional recovery is required, a Tail Gas Treatment Process (TGTU) 

is used. 

FIGURE A-2 TYPICAL CLAUS PLANT BFD20

A.2.3 Plant Capacity 

Table A-9 below shows examples of commercial Claus units at both a large and small scale. 

Treatment capacity sizes range widely; as mentioned above, it is rare to find commercial units 

below 25 TPD. Examples of these types of facilities may have other factors that explain their 

existence, such as age or regulatory issues. It is unlikely that Claus technology would be chosen 

for a grass-roots facility today if the treatment capacity is under 25 TPD. 

Other examples of Claus units built at sizes close to these can be provided, since the technology 

is commercially widespread. 

 

20 Laurance Reid Gas Conditioning Conference, February 2000, courtesy of Charles Kimtantas, Bechtel 

Task 1 Final Report A-6 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Table A-9 Claus Sulfur Recovery Capacities21

Company Location Capacity (TPD, S) 

LARGE 

Duke Energy Gas 

Transmission Co. 

Pine River, B.C., Canada 2,000 

SMALL 

San Joaquin Refining Bakersfield CA, USA 3 

A.3 NATURAL GAS REFORMING 

A.3.1 Introduction 

Natural gas reforming is the process of converting natural gas feedstocks to a synthesis gas 

consisting of CO, CO2, and H2 using steam as a reactant. The synthesis gas will also contain 

small amounts of unreacted methane and residual steam. This process is the first step in many 

industrial processes including methanol synthesis, ammonia manufacture, gas-to-liquids (GTL), 

and hydrogen production. The earliest steam methane reformers were introduced in the U.S. in 

the 1930s by Standard Oil of New Jersey at their Bayway and Baton Rouge refineries for 

hydrogen production from offgases22. 

In the U.S., ammonia and methanol plants, both users of SMR, have been closing as a result of 

competition with foreign plants with access to inexpensive natural gas. Overseas, large “Mega 

Methanol” plants in the Caribbean and in Asia are being built23, along with large GTL plants in 

the Middle East (Qatar)24. Since these projects are all based on SMR, there is expected to be an 

increase in worldwide SMR capacity this decade. Capacity increases of SMR are also expected 

in refineries to produce hydrogen for processing the increasingly heavy, sour crudes being 

brought into petroleum refineries. In the long-term, if the US begins the transition to a 

“hydrogen economy”, SMR will form the bridge providing hydrogen until advanced, nonhydrocarbon source hydrogen becomes feasible. 

Several of the technologies discussed in other sections (Methanol, Ammonia and GTL) depend 

on production of synthesis gas from SMR. The syngas generation step is a very large portion of 

the total cost, on the order of 60% by GTI estimates. Therefore, in order to not unduly bias the 

modular cost study with what substantially amounts to a single technology, it would be advisable 

not to choose more than one technology that has SMR included in it. 

A.3.2 Process Description 

Natural gas reforming is the indirect heating of natural gas and steam in a furnace over a catalyst. 

The catalyst is typically nickel on alumina, and is usually packed within tubes suspended in the 

radiant section of a furnace. Hundreds of tubes are necessary, depending on the capacity of the 

reformer, each 25 to 40 feet in length and 3 to 5 inches in diameter. Sufficient steam must be 

provided to avoid carbon formation. 

 

21 Oil and Gas Journal, June 24, 2002, Worldwide Sulfur Production chart 

22 Catalyst Handbook, Second Edition, Martyn V. Twigg, ed., Wolfe Publishing Ltd., 1989 

23 Gas-to-Liquid News, Methanol Watch, Vol VII, No. 10, October 2004 

24 “In Qatar, Oil Firms Make Huge Bet on Alternative Fuel”, The Wall Street Journal, February 15, 2005 

Task 1 Final Report A-7 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
The reformate consists of CO, H2 and CO2, residual feed methane, and steam. High and low 

temperature shift reactors often follow SMR to adjust the H2 to CO ratio of the product syngas. 

Following the radiant reformer section, there will be a convective zone where heat is recovered 

and steam for the process is produced. Since the catalyst is poisoned by sulfur, the feed is 

usually desulfurized to <1 ppm upstream of the reformer. Reformers are built by many 

companies including, ICI (Johnson Matthey/Synetix), Uhde, KBR. 

The operating pressure is dependent on the desired composition downstream of the reformer, but 

economics usually dictate a range of 280 to 420 psig. The reformer furnace typically runs from 

1400 to 1500 degrees F. Furnaces are typically fired with natural gas. Figure A-3 gives an 

overview of the process25. 

Figure A-3 Steam Methane Reformer Process Flow, Ammonia Plant 

A.3.3 Plant Capacity 

While technological improvements in SMR have been fairly limited, technical learnings, 

economies of scale, and downstream demands have promoted the development of ever-larger 

facilities. Considerable development has gone into mechanical issues such as tube metallurgy 

and schemes to attach the hot tubes to cold headers in order to prolong tube life and avoid 

shutdowns. 

Table A-10 below shows examples of large and small scale SMR facilities. Large scale facilities 

tend to be associated with large chemical plants (ammonia or methanol production), while small 

scale reformers are more associated with hydrogen production. Besides today’s commercial 

small scale plants, natural gas reformers are being built for vehicle hydrogen fueling stations and 

to produce hydrogen for fuel cells. GTI has built a prototype unit to produced 17,000 SCFD of 

hydrogen, although an expected commercial scale unit would be 2 to 2.5 times that capacity. 

This plant requires ~0.01 MMSCFD of natural gas feed. 

 

25 Information from Uhde GmbH 

Task 1 Final Report A-8 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Table A-10 SMR Plant Capacities26

Company Location Capacity (MMSCFD of 

natural gas) 

LARGE 

Burrup Fertilizers Australia 72.6 

SMALL 

Various (fuel cell applications) various 0.1 to 0.01 

A.3.4 Technology Advancements 

KBR has developed the KRES (KBR Reforming Exchanger System) that integrates the offgas 

heat energy from the SMR to perform additional reforming in a specially designed reactor. 

There are two commercial KRES units, both in ammonia plants. The first was started up at 

Methanex's plant in Kitimat, BC, Canada in 1994 and has a capacity of 350 MTD. The second 

started up in 2003 at Liaohe's plant in Panjin, China and has a capacity of 1100 MTD. KRES 

cannot operate as a standalone system, however, but does lower the overall energy demand of 

syngas generation compared to SMR alone. 

Developments in the area of autothermal reforming in lieu of conventional SMR are also in 

progress. 

A.4 METHANOL FROM NATURAL GAS 

A.4.1 Introduction 

Methanol is an important industrial chemical and intermediate for other chemicals including 

silicones, refrigerants, adhesives, specialty plastics, coatings, textiles, and water-treatment 

chemicals. Methanol is produced from syngas that is reacted over a catalyst. The catalytic 

synthesis approach was introduced by BASF in 192327. The primary source of syngas, similar to 

the situation for hydrogen production, ammonia synthesis, and GTL processes, is from reforming 

natural gas. 

A.4.2 Process Description 

The predominant commercial technology is the Low Pressure Methanol synthesis route 

pioneered by ICI (now Johnson Matthey/Synetix). This is a catalytic route using a copper/zinc 

oxide/alumina catalyst. The key design issue when dealing with methanol plants is controlling 

the heat release from the exothermic reactions. This is accomplished by a variety of means 

including placing the catalyst in tubes surrounded by water, interstage cooling, or quenching 

with feed in a multiple-bed reactor. Following this step, crude methanol is refined via 

distillation. 

The manufacture of methanol requires synthesis gas that is usually produced from steam 

methane reforming (see Section 3 for a description of SMR). Following SMR, the synthesis gas 

is compressed to 700 – 1500 psig. The gas leaving the methanol synthesis reactor is cooled to 

 

26 Information on the fertilizer plant from The Fertilizer Institute, http://www.eia.doe.gov/oiaf/aeo/conf/pdf/ vroomen.pdf 27 Catalyst Handbook, Second Edition, Martyn V. Twigg, ed., Wolfe Publishing Ltd., 1989 

Task 1 Final Report A-9 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
condense the crude methanol product, with the remaining gas is recompressed, mixed with fresh 

feed, and returned to the methanol synthesis reactor. The crude methanol is then distilled to 

remove impurities. Figure A-4 gives a schematic representation of a typical methanol process28. 

FIGURE A-4 LOW PRESSURE METHANOL PROCESS FLOW DIAGRAM 

A.4.3 Plant Capacity 

Methanol plants have been increasingly growing in size. The newest facilities, the so-called 

“Mega Methanol” plants, are exemplified by the 1.7 million tonnes per year Atlas Plant in 

Trinidad, which started up in mid-200429. An example of large and small scale commercial 

facilities can be seen in Table A-11. 

Table A-11 Methanol Plant Capacities 

Company Location Capacity (metric tpa) 

LARGE 

Methanex/BP Pt. Lisas, Trinidad 1,700,000 

SMALL 

Praxair Geismar, LA 20,000 

As in many other industries, the trend to larger sizes is to reduce the unit production costs. 

With the availability of low-cost natural gas feedstock in Trinidad, landed methanol prices are 

significantly less expensive than U.S. domestic sources using high-cost natural gas feed. As a 

result, numerous smaller, less efficient plants in the U.S. have closed. In 2003, a number of US 

plants ranging from 80,000 to 200,000 metric ton per year were shut down. With world 

 

28 Information from Methanol Plant Technology – Synetix Low Pressure Methanol (LPM) Process Brochure, Synetix, 2003 

29 More information on this facility can be found at the Methanex website: http://www.methanex.com/ourcompany/locations_trinidad.html

Task 1 Final Report A-10 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
methanol prices low, this trend is expected to continue with even larger plants. Estimates are 

that 4 million tons of annual methanol capacity will be shut down by 200830. Larger plants with 

low-cost transportation options and adequate supplies of low-cost feedstocks will continue to be 

built as supply/demand balance allows. These will largely be Mega Methanol units. 

A.5 AMMONIA PRODUCTION 

A.5.1 Introduction 

Ammonia is produced via the Haber-Bosch process by the reaction of nitrogen and hydrogen. 

The main source of hydrogen in the United States is from steam reforming of natural gas and 

partial oxidation of oil residues. The production of hydrogen by steam methane reforming is 

covered in Section 3. Nitrogen is introduced as compressed air into the secondary reformer. The 

hydrogen reacts with the nitrogen at high temperatures (750-1200oF) and pressures (1500 to 3000 

psig) over a catalyst, usually iron oxide with a potassium hydroxide promoter, although some 

new ruthenium-based catalysts are being suggested. The anhydrous ammonia is sold as product 

or used to produce a variety of fertilizers and other products, including urea, nitric acid, 

ammonium nitrate, and ammonium sulfate31. 

Figure A-5 shows the change in world-scale capacity over time. Prior to the 1970s, plants had 

much greater energy consumption rates, 10-12 Gcal/ton, compared to the current designs of less 

than 7 Gcal/ton. 

FIGURE A-5 AMMONIA PRODUCTION CAPACITY AND ENERGY CONSUMPTION32 (1960-2010) 

 

30 HPI Construction Boxscore, February 2005 

31 Process information from “Energy Use and Energy Intensity of the U.S. Chemical Industry”, Worrell, et. al, Lawrence Berkeley National 

Laboratory, LBNL-44314, April 2000 32 Information from Halliburton, available at http://www.halliburton.com/kbr/relatedInfo/chemicals /modamplt.pdf 

Task 1 Final Report A-11 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
A.5.2 Process Description 

Ammonia is generally manufactured from natural gas using the steam reforming process as seen 

in Figure A-6. While other feedstocks and processes are used, this process represents the 

majority of the commercial designs seen today. There are several reaction stages and catalyst 

reactions key to the economic operation of modern ammonia production plants. 

The first stage is purification where impurities, mainly sulfur compounds, are removed from the 

gas stream. Steam reforming is performed in two stages. In the primary stage, the endothermic 

reactions take place at pressures around 450 psig and temperatures of 1450°F or higher. This is 

followed by an exothermic secondary reformer, where air is added to the partially reformed gas 

stream. 

The carbon monoxide in the gas leaving the secondary reformer is converted to carbon dioxide in 

the shift reactors. The carbon dioxide is then removed by scrubbing the gas stream via an amine 

process. Any residual carbon oxides are then converted back to methane before compression to 

ammonia synthesis pressure. The final reaction stage is ammonia synthesis where the hydrogen 

and nitrogen combine to form ammonia. This reaction stage takes place at high pressure (1500 to 

3000 psig) and is highly exothermic. The removal of product ammonia is accomplished via 

mechanical refrigeration or absorption/distillation. Typically, refrigeration is more economic 

for high pressure operation, while absorption/distillation is favored for lower pressure operation. 

FIGURE A-6 AMMONIA PRODUCTION FLOWSHEET33 

A.5.3 Plant Capacity 

Major modern plants can easily produce 1,000 tonnes of ammonia per day, requiring almost 1 

million m3 of methane. Due to this high feedstock requirement and large plant economies of 

scale, production is moving to low cost hydrocarbon sources. Alternative sources of hydrogen 

 

33 Information obtained from Synetix, Introduction to Ammonia Manufacturing, available at 

http://www.synetix.com/ammonia/pdfs/ammonia.pdf 

Task 1 Final Report A-12 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
are being sought, including coal via gasification and water via electrolysis using an alternative 

energy source, such as solar. Many ammonia plants in the United States are currently shut down 

because the high cost of feedstock results in production costs greater than the market value of the 

produced ammonia34. 

Table A-12 Ammonia Plant Capacities 

Company Location Capacity (TPD) 

LARGE 

Burrup Fertilizers Australia 2,200 

SMALL 

Shoreline Chemical Gordon, GA 31 

A.6 GAS-TO-LIQUIDS 

A.6.1 Introduction and Summary 

GTL or Gas-to-Liquids, represents a number of processes which have been proposed to monetize 

stranded natural gas by manufacturing mainly low sulfur diesel and side products like waxes and 

lube oils from low-cost natural gas. Production technology is essentially synthesis gas 

manufacture by reforming, followed by proprietary synthesis technology based on FischerTropsch catalytic chemistry. 

Only two commercial plants exist: the Shell Bintulu facility in Malaysia, and Sasol’s commercial 

technologies in South Africa. All others facilities are either pilot units or are currently in the 

design stages. A number of process technologies exist for conversion of natural gas to FischerTropsch liquids. All GTL technology is highly proprietary, making it difficult to provide 

accurate process and cost information. With the potential for wider commercial emergence of the 

technology in the future, the GTL market is highly competitive unlike semi-public domain 

technology like SMR and methanol production. While small scale GTL test facilities exist, costs 

for pilot plants are roughly 3 times as much as they would be for a scaled-down unit of same size 

based on commercial technology35. Costs for pilot scale facilities can also be deceptively low 

because, for example, they may not include the product upgrading step or do not require a 

commercial oxygen facility for a unit of that size. 

Because of the difficulties in developing accurate cost and design information for GTL plants 

with published data, it is not recommended that they be considered for future design work. 

A.7 NATURAL GAS LIQUIDS EXPANDERS 

A.7.1 Introduction 

Natural gas that is extracted from terrestrial sources consists predominantly of methane, and, to a 

lesser extent, a quantity of Natural Gas Liquids (NGL). These are typically ethane and propane, 

but also contain some butane and lesser amounts of higher carbon content, aliphatic 

 

34 US Geological Survey (USGS), available at http://pubs.usgs.gov/of/2004/1290/). 35 Based on GTI estimates and Merrow, Phillips, and Myers, Understanding Cost Growth and Performance Shortfalls in Pioneer Process 

Plants, The RAND Corporation, 1981 

Task 1 Final Report A-13 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
hydrocarbons. Table A-13 below shows compositions for a range of potential feedstocks to a 

NGL plant. 

Table A-13 Typical NGL Feedstock Compositions 

Classification 

Super-Rich Rich Lean Super-Lean 

C1 (mole%) 82.00 87.00 92.00 96.30 

C2 9.70 7.90 4.80 2.20 

C3 4.50 3.30 1.80 0.80 

IC4 0.50 0.46 0.49 0.20 

NC4 1.80 1.13 0.87 0.50 

IC5 0.50 0.00 0.00 0.00 

NC5 0.49 0.21 0.04 0.00 

C6+ 0.51 0.00 0.00 0.00 

BTU/scf 1251.81 1161.77 1105.35 1054.51 

If the quantities of these materials are low, they slightly enhance the heating value of the natural 

gas and may be left in the gas. If the amount of such materials is such that there may be 

condensation in the transmission pipeline, or there could be an impact on the suitability of the 

gas for various applications (usually measured in terms of the heating value and/or the related 

Wobbe number), the pipelines may not purchase the gas from producers without prior treatment 

to remove the heavy hydrocarbons. 

In addition, a large industry has grown to utilize the ethane/propane and heavier hydrocarbons 

being removed from the natural gas as petrochemical feedstocks. Extraction of NGLs from 

natural gas can be either profitable or an unavoidable cost of production of natural gas for 

pipeline transport, depending on the relative price of the hydrocarbons on a heating value and 

chemical feedstock basis. 

A.7.2 Process Description 

Early facilities used circulating heavy hydrocarbon liquids to absorb the lighter components in 

lean oil absorber plants. In the 1960s, advances in turbomachinery led to newer cryogenic 

processes being introduced. These turboexpander processes have become the dominant 

technology. 

Technology has evolved mainly to enable the expander plants to have reasonable recoveries and 

economics when rejecting ethane (leaving it in the gas) while recovering 90% or more of the 

propane and heavier components. This is necessary to enable such plants to remain in operation 

when economics favor leaving the ethane in the pipeline gas. There are a number of patented 

schemes that improve upon the original turboexpander design in order to enable this mode of 

operation. Many companies offer NGL expander plants including Fluor, Ortloff, IPSI LLC, 

ProQuip, and ABB Randall. 

Task 1 Final Report A-14 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
Figure A-7 shows the schematic of the NEPTUNE II NGL extraction plant built and operated by 

Enterprise at St. Mary Parish, LA. This plant uses technology from IPSI LLC, a subsidiary of 

Bechtel, and has a gas processing capacity of 300 MMSCFD. The NGL content of the inlet gas is 

about 4.0 gallons per MSCF. Approximately 88% of C2, 99% C3 and 100% of C4+ are 

recovered as NGLs. The shrinkage, or loss of feed natural gas resulting from the processing, is 

about 11-12% of the inlet gas throughput. This plant was commissioned in early 2004. 

The process for the removal of NGL from natural gas is essentially a mechanical refrigeration 

scheme. There are no chemical reactions taking place. Natural gas is fed at a pressure above 

400 psi, is exchanged with exiting cold gas from the top of the demethanizer column, then is sent 

to a cold separator to drop out condensed liquids. The condensed liquids are further cooled and 

sent to the demethanizer column as cold reflux. The gas exiting the cold separator is expanded in 

the turboexpander that cools and condenses significant amounts of liquid for column reflux. The 

demethanizer distills the cold liquid and gas feeds to the column. The overhead gas is highly 

concentrated in methane with some ethane if desired. The column bottoms will be primarily 

ethane and propane. Depending on the gas composition and the NGL product specifications, it 

may be necessary in some plants to further distill the liquid to remove butane and heavier 

components. 

FIGURE A-7 SCHEMATIC FOR NEPTUNE II TURBOEXPANDER PLANT36

 

36 Oil and Gas Journal, July 21, 2003 

Task 1 Final Report A-15 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
A.7.3 Plant Capacity 

Plants are constructed of a size commensurate with the amount of NGL expected in the natural 

gas feed. Natural gas contains NGLs at varying concentrations, and different producing regions 

may have varying amounts of natural gas being produced depending on the age of the field and 

the number of wells being drilled. A fairly wide range of expander plants sizes are currently 

being built. This is unlike other process technologies mentioned during this study, where 

economies of scale have led to increasingly larger facilities. 

The number of expander plants in the U.S. is on the order of 300 – 400. The largest facilities 

have capacities on the order of 2,500 MMSCFD, with multiple trains of processing. Smaller, 

skid-mounted plants may be as small as 10 to 50 MMSCFD. Table A-14 gives examples of 

typical large and small scale NGL Expander facilities. 

Table A-14 NGL Turboexpander Plant Capacities 

Company Location Capacity (MMSCFD) 

LARGE 

Enterprise St. Mary Parish, LA 300 - 350 

SMALL 

Aqila Sommerville, TX 25 

A.8 CRYOGENIC OXYGEN PLANTS 

A.8.1 Introduction 

Oxygen was first extracted from the atmosphere by a chemical process. This was superseded in 

the 1920s by the cryogenic process involving low temperature liquefaction and distillation of 

air37. Over the years, the sizes of the oxygen plants have grown from 100 TPD to almost 4,000 

TPD. The cryogenic air separation process is by far the most widely used. However, noncryogenic techniques first developed during the 1970s – adsorption technologies – are becoming 

increasingly significant for smaller or lower purity applications. 

The choice of oxygen production technologies depends on several factors: the volume required, 

the purity required, the customer location, and the oxygen usage pattern. For the large tonnage 

oxygen customers (over roughly 300 TPD), a cryogenic plant is usually more economic. For the 

customers requiring lower oxygen flow rates, adsorption or delivered oxygen is usually 

preferred. Cryogenic plants cannot compete in the low tonnage area primarily because of the 

difficulty of scaling down the distillation column and the poor efficiencies of the small 

compressors. Non-cryogenic plants, on the other hand, are limited in their sizes because of the 

constraints of the adsorber diameters. 

The major oxygen suppliers are Praxair, Air Product, BOC Gases (UK), Air Liquide (France), 

and Linde (Germany). Hundreds of cryogenic and non-cryogenic plants have been installed 

worldwide. 

 

37 Information on technology history from BOC gases, available at http://www.boc.com/capability/gases /air_separation.asp 

Task 1 Final Report A-16 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
A.8.2 Process Description 

In the cryogenic process, filtered air is first compressed and cooled by the main air compressor 

and is sent to the adsorbers, where water, carbon dioxide and any hydrocarbons are removed 

from the air stream. The clean dry air is sent to a series of heat exchange and distillation steps to 

separate the major constituents: oxygen, nitrogen, and argon. 

FIGURE A-8 GENERIC CRYOGENIC OXYGEN PLANT FLOW DIAGRAM38

A.8.3 PLANT CAPACITY 

As with many other process technologies discussed earlier, cryogenic oxygen plants do have an 

economy of scale that has led to larger and larger plants in recent years. Most small facilities 

either are older, or have specialty niche applications that justify their use. Table A-15 below 

shows examples of small and large scale cryogenic oxygen facilities39. 

Table A-15 Cryogenic Oxygen Plant Capacities 

Company Location Capacity LOX (TPD) 

LARGE 

Air Liquide Secunda, South Africa 3,550 

SMALL 

BOC Lahore, Pakistan 30 

 

38 From Air Products, available at http://www.airproducts.com/NR/rdonlyres/C13A9654-7BB1-4337-A4D9-94F64546A672/0/APACK.PDF 39 Air Liquide information from Engineering News, available at http://www.engineeringnews.co.za/eng /features/sasol. BOC 

information from “First Move of Aspen Plant Ends in Success”, FrostByte – Newsletter of Cryogenic Industries, Vol. 10, No. 4, Spring 

2000 

Task 1 Final Report A-17 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
A.9 ETHYLENE AND OTHER OLEFINS PRODUCTION 

A.9.1 Introduction 

Ethylene is used primarily as an intermediate in the manufacturing of plastics and other 

chemicals. It is polymerized to form polyethylene, chlorinated to form the precursors for 

polyvinyl chloride, or reacted with benzene in the manufacturing of polystyrene. Other chemical 

uses include the production of ethylene oxide, ethanol, polyvinyl acetate, and ethylene glycol. It 

also has a small market as a plant hormone for ripening picked fruit. 

The primary route to the production of ethylene is by thermal cracking of hydrocarbons. Other 

paths, including catalytic pyrolysis, oxy-dehydrogenation of ethane, oxidative coupling of 

methane, and dehydration of ethanol have limited or no commercial interest. In thermal 

cracking, gaseous or light hydrocarbons are briefly heated to 1500-2000oF, initiating a free 

radical reaction, breaking down the heavier hydrocarbon chains into smaller ones, and 

unsaturating previously saturated hydrocarbons. The end product is a complex mixture of 

ethylene, propylene, butadiene, benzene, fuel oil, and other products, requiring extensive 

separations and purification steps. 

A.9.2 Process Description 

The heart of the ethylene plant is the cracking furnaces. These are tubular reactors in which the 

pyrolysis reactions take place. Before 1960s, the cracking tubes were horizontal rows in a 

radiant chamber. This gave low ethylene yields, leading to more modern designs using vertical 

tube rows to provide superior mechanical performance and higher capacity. Steam is added to 

the feed to prevent coking and to minimize side reactions. The product mix ranges from 

hydrogen to tars. This requires extensive separations to meet product purity specifications. This 

is performed by compression, drying, fractionation, absorption, adsorption and catalytic 

conversions. 

A simple flowsheet for an ethylene production process is shown in Figure A-940. Figure A-9 

shows a plant with a relatively light feedstock mix, requiring little separation of hydrocarbons 

heavier than propane. The complexity of the flowsheet will depend on the feedstock and product 

mix desired. 

Feedstock enters the plant and is fed to the pyrolysis heaters after preliminary heat exchange. 

Here the feedstock is broken down to lighter hydrocarbons. The remainder of the plant is 

separations and purification. The gasoline and fuel oil fractions are recovered first through 

quench and stripping that removes the heaviest fractions. The vapor fraction is then compressed 

to over 500 psig for additional purification via fractionation and cold box steps. Acid gas 

removal may be performed between compression stages41. 

 

40 Information from Halliburton, available at http://www.halliburton.com/kbr/relatedInfo/chemicals/ ethylene.pdf 41 Technical information from Encyclopedia of Chemical Technology, Fourth Edition, Kirk-Othmer, John Wiley & Sons, NY, 1994 

Task 1 Final Report A-18 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
FIGURE A-9 KBR SCORETM (SELECTIVE CRACKING OPTIMUM RECOVERY) ETHYLENE 

TECHNOLOGY 

A.9.3 Plant Capacity 

The capacity changes over time are shown in Figure A-1042. The solid line is an estimate of the 

maximum plant capacity able to be built during that year. Through the 1960s and 1970s, steam 

cracking focused on gaining economies of scale by dramatically increasing the size of each plant. 

The increase was facilitated by the evolution of compression equipment to highly reliable, large 

centrifugal compressors, improvements in furnace design, a better understanding of light ends 

fouling, and advances in process control. Also significant were improved computerized 

simulations that made design of today's modern recovery equipment possible. 

 

42 From R. Orriss, “Major Investment Opportunities in the 21st Century, The Changing World of the Global Ethylene Business,” Global 

Ethylene Conference, Saudi Arabia, 2004. 

Task 1 Final Report A-19 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
FIGURE A-10 ETHYLENE PLANT SIZE PROGRESSION, 1945-2005 

The drive to build larger single train steam crackers slowed after about 1980. This was not 

caused by technological barriers, but by the difficulty of marketing large quantities of product 

and acquiring enough feedstock. Focus shifted to debottlenecking existing units. The 

ExxonMobil plant at Baytown, Texas, is currently the largest single train in the world at about 

1,200,000 tons/yr. As the dotted line shows, this plant was built in 1979 with an original 

capacity of 590,000 tons/yr. Through several debottleneck programs, the capacity has been 

doubled at far below grassroots cost. With improved heat transfer and mass flow inside the 

furnace tubes, better process models, extremely short times in the reaction zone, and rapid 

quench systems, the world-class plants have the capacity to expand from 600,000-1,000,000 

tons/year to over 1,500,000 tons/y. 

Table A-16 shows examples of large and small ethylene plants currently in operation. Small 

scale facilities usually have some sort of feedstock advantage that justifies their continued 

operation. For example, niche facilities with low-cost liquids that have superior economics for 

ethylene production over blending to gasoline or other refinery processes can economically exist. 

Table A-16 Ethylene Plant Capacities43

Company Location Capacity (TPD) 

LARGE 

ExxonMobil Baytown, TX 3,500 

SMALL 

Javelina Corpus Christi, TX 108 

 

43 Information on plant sizes from team personal experience, Oil & Gas Journal, March 22, 2004, and Chemical Markets Reporter, October 6, 

2003 

Task 1 Final Report A-20 

Cost Estimates of Small Modular Systems 

United States Department of Energy/National Renewable Energy Laboratory 
F1146-E(12/2004) 

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DE-AC36-99-GO10337 

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4. TITLE AND SUBTITLE

Equipment Design and Cost Estimation for Small Modular Biomass 

Systems, Synthesis Gas Cleanup, and Oxygen Separation 

Equipment; Task 1: Cost Estimates of Small Modular Systems 

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NREL/SR-510-39943 

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BB06.3710 

6. AUTHOR(S)

Nexant Inc. 

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San Francisco, CA 94105 

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ACO-5-44027-00 

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This deliverable is the Final Report for Task 1, Cost Estimates of Small Modular Systems, as part of NREL Award 

ACO-5-44027, “Equipment Design and Cost Estimation for Small Modular Biomass Systems, Synthesis Gas Cleanup 

and Oxygen Separation Equipment”. Subtask 1.1 looked into processes and technologies that have been 

commercially built at both large and small scales, with three technologies, Fluidized Catalytic Cracking (FCC) of 

refinery gas oil, Steam Methane Reforming (SMR) of Natural Gas, and Natural Gas Liquids (NGL) Expanders, 

chosen for further investigation. These technologies were chosen due to their applicability relative to other 

technologies being considered by NREL for future commercial applications, such as indirect gasification and fluidized 

bed tar cracking. Research in this subject is driven by an interest in the impact that scaling has on the cost and 

major process unit designs for commercial technologies. Conclusions from the evaluations performed could be 

applied to other technologies being considered for modular or skid-mounted applications. 

15. SUBJECT TERMS

biomass; Nexant; small modular; synthesis gas; syngas; fluidized catalytic cracking; steam methane reforming; 

natural gas liquids expanders; scaling; National Renewable Energy Laboratory; NREL 

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