Mega-olefin plant design: Reality now
Bigger is better, when planning future ethylene processing capacity
C. P. Bowen and D. F. Jones, Shaw Energy & Chemicals, Houston, Texas
The olefin industry continues to expand and upgrade its capacity and equipment. Present olefin-
plant capacities, investment costs, equipment delivery schedules and construction timetables
represent many of the industry's major challenges. These concerns were true in the late 1990s
when the largest cracking furnaces (ethane and naphtha types) were constructed, and when the
largest total plant capacities approached 1,500,000 tpy (1.5 MMtpy).
As current projects, prices and participants have extended plant sizes and
upstream/downstream links, these challenges must be addressed in conjunction with key
equipment manufacturers. Ongoing design of key equipment corresponds to respective nominal
naphtha/ethane cracking capacities of 1.5 MMtpy to
2 MMtpy for future ethylene plant designs.
Market trends. Olefin production capacity has virtually doubled in the past 15 years. Current
capacity is approximately 130 MMtpy and demand is about 114 MMtpy. Global ethylene growth
rate is still approximately 4.5%/yr, and processing capacity is forecast to double again by 2025.
Various alternative production routes to ethylene/propylene have been recently proposed, but
the bulk of new production will still come from conventional steam crackers.
Corresponding plant capacity projections. As a consequence to the total olefin growth rate
during the past generation, steam cracker maximum capacities have increased from 0.5 MMtpy
ethylene (late 1970s) to 1.5 MMtpy today. At present, there are over 30 units with 1+ MMtpy of
processing capability in service or under construction. The industry has correspondingly
increased individual equipment item sizes, and, in anticipation of even larger plant capacities, is
now developing even larger key equipment items.
The steam-cracking industry is not unique in this trend. Refinery units have doubled during the
same time frame (200 Mbpd to 400 Mbpd); petrochemical derivative units have also increased
unit size by a similar factor as per ethylene units.
Investment cost vs. capacity. A standard relationship between investment and plant capacity
for most olefin units is linked to a factor ranging from 0.65 to 0.85. Accordingly, larger olefin
units can generate higher cost production margins. But this involves several investment
concerns including equipment and operational reliability, availability of essential feedstocks and
adequate product disposal.
Feedstock, energy and chemical plant investment costs have significantly increased during the
last five years. At the turn of the century, the reference cost for a standard naphtha-cracking
ethylene plant was ±$1,000/ton ethylene per year. At present, that cost is 40% higher due to the
higher cost of basic and specialty materials, increasing fabrication sales costs, construction
team availability and general engineering costs. So, the combination of higher unit costs offset
by investment costs vs. capacity factors further encourages the olefins industry to move to
higher capacity facilities.
Fig. 1 Current mega-plants incorporate significantly higher
cracking furnace capacities and performance flexibility.
Equipment size trends. The industry has responded well to the increasing plant capacity
demands and improved operational efficiency. Accordingly, equipment disposition and item
count within today's olefin units is actually less than it was a generation before. New plant
designs have increased specific capacities of key items significantly. Newer olefin facilities have
integrated several heat and mass transfer facilities, combined various equipment vessels,
expanded catalytic reactor capacities, extended heat pump links between standard
compressors and fractionators, and significantly increased key equipment duties/capacities.
Specific data related to key olefin equipment items will be further discussed.
Cracking furnaces. Early smaller ethylene plants often contained up to 20 individual cracking
furnaces. For example, a major European unit, constructed 35 years ago, had an initial
processing capacity of 450 Mtpy and used 16 fresh-feed furnaces plus two recycle furnaces.
Today, the standard liquid-feed furnace sizes are seven times larger. Gas-feed cracking
furnaces produce even higher ethylene yield outputs.
The latest designs for liquids-cracking furnaces use symmetrical U-type radiant coils; thus,
furnace capacities have increased by up to 25% within identical radiant firebox dimensions. One
of the largest naphtha-cracking furnaces currently under construction will produce 190 Mtpy of
ethylene. Key design features in this unit are:
• Dual radiant cell design (each cell contains 112 U-type coils)
• Total radiant firebox floor firing with relatively large low-NOx burners
• Radiant cells may be separately decoked; decoking effluent directed into radiant firebox
to minimize decoke effluent contaminants
• Multi-pass convection section including slot for future SCR—selective NOx catalytic
• Variable RPM induced flue gas fan to optimize combustion conditions during
• Modular system design-fabrication of entire furnace to facilitate construction phase
• Individual double-tube selective linear exchangers (SLEs) are directly connected to
each radiant coil outlet; total quench heat recovery is achieved in each SLE per each
• For heavy feedstocks, the cracked effluent is cooled to at higher temperature and is
subsequently directly quenched with quench oil to reduce main transfer line operating
• Multi-stage startup/shutdown control facilities provide necessary uniform operations and
avoid equipment damage due to emergency conditions
• The 190-Mtpy U furnace consumes about 200 MW fuel and generates 100 metric
tons/hr super high-pressure (SHP) steam
• Operating run length for such furnaces is typically about 45 days.
Representative cracking furnace capacities are listed in Fig. 2.
Fig. 2 Recent North American mega-furnace designs.
Main train compressors. Virtually all ethylene plants today incorporate a single cracked-gas
compression system. Earlier units sometimes used twin parallel units or even a spare unit.
Some units prior to 1970 consisted of multiple reciprocating compressors. High-capacity
centrifugal compressors represent many design, fabrication and operational challenges. Today,
there are few compressor manufacturers who can meet mega-capacity requirements. Recent
design extensions and definitions include:
• Naphtha-type cracking requires low suction pressures (0.17 barg– 0.30 barg) to
optimize cracking selectivity. Discharge pressure is typically about 35 barg to permit
downstream condensation of methane vs. ethylene refrigerant. Maximum interstage
discharge temperatures should ideally not exceed 90°C. Accordingly, such system
normally requires five stages in series.
• Typical suction volumetric flow for a 1.5-MMtpy ethylene naphtha cracker compressor
ranges from 500,000 m3/h to 700,000 m /h. The LP body for recent larger scale plants
has either been a large, double flow machine, or, in some plants two parallel flow LP
bodies. Typically, the impellers in process Stage 2 (MP body section) are very similar in
flow coefficient to each of the double flow impellers in process Stage 1 (double flow LP
body or two parallel single flow bodies). Each stage often consists of three impellers.
But the diameter and rotating shaft length for the 3D impellers sometimes limit the
impeller count for such double flow casings to only two impellers per stage.
• The cracked-gas compressor (CGC) will either consist of three or four casings,
depending on capacity and vendor. The three casing arrangement is normally a double
flow LP casing for stage one; a single flow MP casing for Stages 2 and 3; and a single
flow HP casing for Stages 4 and 5. In some cases, vendors have selected four
compressor bodies, two parallel flow LP bodies for Stage 1 and the MP and HP bodies
similar to the three body arrangement described previously.
• Client input is required during the selection of the first-stage compressor flow
coefficient— = f (volume flow, diameter , RPM). Long-term plans to increase plant
capacity will impact this decision (See Fig. 3).
• Client input is also required for train layout considerations. For 1.5-MMtpy-naphtha
crackers, the cracked gas API normal condition can absorb power in the range of 90
MW to 100 MW. Operating two trains in series, LP/MP and HP, respectively, is to be
considered. In a single-train arrangement, normally the impeller/diffuser pairs selected
for Stages 1 and 2 will operate at optimum efficiency. The performance for Stages 3, 4
and 5 are progressively compromised by the rotational speed dictated by Stages 1 and
2. In mega-plant sizes, the power saved by operating Stages 4 and 5 at a higher
optimum speed can be significant.
• Ethane-cracking compressors can be optimally designed for higher suction pressures
(~1.3 barg) and slightly lower discharge pressures, and accordingly have only four
stages. Ethane cracking yield molecular count is about two whereas naphtha yield is
approximately four molecules per feed molecule. Ethane-cracker cryogenic fractionation
is essentially ethylene/hydrogen; hence, the demethanizer system pressure is
somewhat lower than that of the naphtha generated ethylene/methane/hydrogen
system. Accordingly, the ethane-cracker CGC and its upstream quench area system
are smaller in terms of both volumetric flow and energy demand.
• For reference purposes, ethane-cracker CGC capacity and power are typically only
75% of that of the equivalent ethylene capacity naphtha cracker. Based on present
commercially available designs, the corresponding single compressor/driver maximum
capacities are about 1.5-MMtpy ethylene from naphtha and 2 MMtpy ethylene from
• Client input into the final selection of impeller flow coefficient will impact future uprate
potential. It also impacts capital investment (casing size), operating cost (polytropic
efficiency at operating flow coefficient relative to peak efficiency flow coefficient) and
reliability (impeller stress and/or critical speed margin). For example, very high flow
coefficient impellers, bordering on mixed-flow designs, have a significantly higher
aspect ratio (axial stage length vs. diameter) than lower flow coefficient designs. Critical
speed margin can be compromised by selecting very high flow coefficient impellers.
Such high-flow coefficient impellers can also limit train speed based on operating
• Since cracked gas invariably contains concentrations of hydrogen sulfide up to 1,000
ppm, the operating stress limits of the compressor impellers must be lower than that of
the equivalent sulfur-free compression condition. Today, essentially all CGC
compressors are designed per NACE.
• All critical duty centrifugal compressors are equipped with dry-gas seals. Dry-gas seal
systems improve unit reliability and reduce maintenance, such as the avoidance of
lube-oil contamination of the cracked gas and refrigerant inventories associated with
legacy oil seal systems.
• Continuous clean-water injection facilities are installed for CGC compressor
temperature suppression. The vendors have specific limitations on injection rate and
water particulate size to minimize erosion while minimizing temperature rise and its
resultant positive impact on run length. Many users also install an intermittent wash-oil
injection for cleaning. Considering the modest capital required, it is recommended to
design and install both systems during construction. As noted, water injection is
continuous, except when the oil is injected for cleaning. The frequency for cleaning is
site specific and established during operation of the plant.
• Internal coatings of both stationary and rotating elements of the CGC compressor are
used by many olefin facilities as a preventive measure to fouling. Such coatings are
often re-installed at each plant scheduled turn-around.
• An essential element in verifying the integrity of the compressor system is shop testing,
as shown in Fig. 4. Component material verification, testing and dimensional
compliance are consistent with current world-scale plant practices. As with the present
practice, system energy levels demand section by section compressor thermodynamic
performance testing and individual body mechanical testing. The requirement for string
testing is user specific.
Fig. 3 Typical centrifugal compressor impeller configuration
and efficiency vs. flow coefficient, typical trend for a
Fig. 4 Main train compressor testing at the fabrication shop.
Refrigerant compressors. In ethylene plant design, the respective power demands of the
ethylene refrigerant compressor and propylene (or propane) refrigerant compressor are
approximately 1/3 and 1/2 of that of the CGC:
• The C2, C3 refrigerant compressor capacities are smaller than that of the cracked-gas
machine. Ethylene product export is typically high-pressure ambient vapor. When
ethylene users are not available, the plant is often designed for storage, where the
product must be exported temporarily as cryogenic liquid. The vapor product condition
typically defines the API "Normal" condition for the refrigerant compressors. The
intermittent liquid condition often defines the API "Rated" condition. The volumetric
differential between these two conditions can be significant, and can drive the API
Normal condition toward surge. Initial compressor impeller/diffuser selection, number of
stages (head per stage) and relative Mach number (rotational speed) all impact the
compressor stability, and should be evaluated carefully prior to order.
• Depending on the precise recovery scheme design, liquid-ethane-feed cracker C3
refrigerant compressor may have only two or three stages. Naphtha-cracker designs
are typically more complex (~ four stages) due to the minimum feedstock refrigeration
credit and range of products and co-products.
• Most C2 refrigerant compressors, which are typically combined with C2 fractionation
heat pump duty are designed with four stages. In some cases, the C2 refrigerant
compressor duty requires five stages, and two compressor bodies.
• The refrigerant compressors are not exposed to component sulfur; therefore, they are
not required to be manufactured to NACE requirements. However, the relatively low
operating temperatures for these machines require materials of construction that can
maintain sufficient ductility throughout the operating range.
• Depending on the plant location and available utilities, some C3 refrigerant compressors
use air-cooled C3 condensers. The discharge pressure must be correspondingly higher
than water-cooled condenser systems. Final discharge pressure has a second order
impact on the selection, and virtually all C3 systems are designed with a single-body C3
• Certain of the critical sized equipment related to the refrigerant compression systems
are the respective low-temperature plate-fin reboilers/condensers and the refrigerant
condensers (Fig. 5).
Fig. 5 Axially split, radial flow inline compressor for
olefins refrigeration service.
Steam turbine drivers. Key design criteria include:
• Fixed-rpm steam turbine drivers above 200 MW are currently available. Variable-rpm
units, as implemented for ethylene plants, are currently limited to 80 MW. We anticipate
that the next-generation steam turbine variable-rpm drivers will shortly be rated at 100-
• Critical-duty drivers are supplied by steam pressures from 100 barg to 300 barg.
Ethylene plants are designed to supply steam from 66 barg to 130 barg. Higher
pressure systems are typically more efficient; their steam system operating conditions
require very high-quality boiler feedwater (BFW) pre-treatment contaminant removal,
de-ionization and de-phlegmation. High-power turbine drivers incorporate various
intermediate steam pressure extraction systems. Ethylene plant turbines driving
compressors are typically designed as extraction condensing. Normal operating
conditions dictate the extraction section design, and startup normally defines the
condensing section design.
• Close-cycle ethylene plant steam systems involve relatively high capacity steam turbine
condensation fractions. Accordingly, the lowest pressure turbine blades are relatively
long and operating speeds often approach blade mechanical design limits. More
efficient systems, particularly for high-capacity units, maximize medium/low pressure
steam export opportunities and consequently reduce the condensing phase traffic.
Consequently, this design reduces the condensing section volume and required blade
length, and resultant tip speed and mechanical stress.
• Another operating adjustment involves steam turbine condenser exhaust pressure.
During startup or reduced duty operation, the compressor power demand is reduced
and, accordingly, the steam exhaust pressure can be higher. The condensation
temperature is consequently higher; the flowrate can be higher, and the HP or MP
extraction rates can be conveniently reduced. Hence, the turbine blade mechanical tip
speed limits can be consequently avoided.
• It is expected that the mega-plant cracked gas and C3 drivers will both operate on the
highest pressure steam available within the system. The C2 driver can be designed to
operate at the highest (SHP) steam pressure, or possibly an intermediate pressure
(HP), depending on location specifics.
• Over the past two decades, advanced steam turbine coatings have been successfully
installed in olefin plant applications. The resulting reliability and run-length
improvements attributed to steam turbine stationary and rotating coatings warrant
evaluation of these options.
• For the 1.5-MMtpy-naphtha crackers or 2-MMtpy-gas crackers referenced, the CGC
driver requirement is nominally 100-MW. This condition exceeds the power of steam
turbine mechanical drive systems currently operating in the market. Some vendors have
completed feasibility and are willing to build 100 MW, single-driver CGC drivers. As
noted earlier, compressor power consumption optimization compels the consideration of
two drivers. Many other technical factors influence the driver selection including: trip
and throttle valve number and size, valve chest references, number of control sections,
blade references, etc. Preferred vendor selection, user preference and risk aversion will
impact the final decision for one driver vs. two drivers for the cracked-gas train.
Main fractionating towers. Mega-olefin plants include several major towers. Earlier designs
often involved semi-tower delivery and in-situ upper/lower section welding. Today, entire towers
can be delivered without such construction challenges. A 1-MMtpy naphtha cracker ethylene
unit will typically include up to four towers weighing over 1,000 tons. However, freight transport
systems adjacent to ports can handle towers of up to 2,000 tons; equivalent major erection
facilities can also handle such items.
As plants become larger, the vessel diameters increase on a square root ratio. In certain cases,
the diameters and corresponding weight of vessels and towers can be reduced by reducing the
reflux ratio and increasing the tray count. Such optimum dimensions apply to the major ethylene
and propylene fractionators. Other multi-tower sequences can be adapted to reduce duty, e.g.,
combination rectifier/stripper towers.
The largest single duty in most naphtha crackers is the C3 fractionator. Its standard reflux ratio
may be up to 15:1; thus, both tray count and diameter are extended duties. For reference, a
600-Mtpy polymer-grade propylene fractionator will typically consist of a two series tower
system; the respective diameters could be up to 7 m, with 115 m in length and the
corresponding weight up to 1,350 tons.
Another major tower is the quench-oil tower. A heavy-naphtha feedstock 1-MMtpy ethylene unit
quench tower would be approximately 13 m in diameter and up to 45 m tall. Erection weight for
such a unit would exceed 1,100 tons. Fig. 6 shows an example unit.
Fig. 6 The 1,000+ ton quench-oil tower is a
key element in this heavy cracking
feedstock mega unit.
Ethane or ethane/propane crackers incorporate a quench-water tower in a similar role. That
vessel may be of equivalent size/weight to the naphtha-cracker quench-oil tower. But, as
mentioned previously, if the ethane cracking coil outlet pressure is elevated (e.g., 1.8 barg vs.
0.6 barg) the downstream quench system and cracked-gas compression system are of reduced
physical size. Such a high-pressure ethane cracker producing 1.5 MMtpy of ethylene would
incorporate a quench-water tower of only 8.5 m in diameter and 30 m in height with a weight of
approximately 500 tons. If however the same plant was based on lower conventional coil outlet
pressure, the volumetric cracked gas flowrate would be approximately doubled with a
corresponding increase in quench tower diameter and weight.
The other major fractionator tower is the ethylene-ethane splitter. The present current recovery
scheme design minimizes the duty on this tower but it is nonetheless of major proportions and
weight. A 1.5-MMtpy ethylene fractionator is approximately 7 m in diameter and 60 m in length
with a weight of 500 tons.
Major heat exchangers. The largest heat exchangers on many plants are the various air
coolers, which can represent up to 50% of the total plant cooling duty. The installed area of
such mega plants poses problems. In most cases, the horizontal air coolers will be installed
above mega pipe racks—up to 20 m in elevation. Major air coolers will consist of 4–6 rows of up
to possibly 20 m in length. Many Middle East plants have restricted seawater usage for cooling,
but the ambient-air temperature range forces air coolers to operate at high condensing
temperature/pressure. Steam turbine condensers may also be air coolers; these are frequently
installed in triangular overhead formats to maximize available duty per unit plot area.
Shell-and-tube heat exchangers have also been extended beyond their traditional dimensions.
One of the latest mega plants have incorporated single mega exchangers in high-capacity
services such as cracked-gas intercoolers/aftercoolers, refrigerant condensers and fractionator
condensers. Such exchangers are up to 2,000 m surface area per shell with single pass tubes
of 13 m in length. Mega units typically relax the delta in temperatures and pressure drops
across such exchangers and thereby reduce the surface area and process-side flow
dimensions. Key benefits from such designs are linked to the associated piping and installation
Plate-fin exchangers have also been significantly increased in sizes and capacity as a
consequence of mega unit decisions. The manufacturing facilities and techniques have been
upgraded so that unit exchangers volumetric size may extend up to 18 m per core and
parallel/series requirements are minimized. Various related improvements have also been
incorporated into plate-fin exchangers that directly benefit mega-plant systems, e.g., anti-
corrosion metallurgy, upstream protective guard beds, integrated tower reboilers/condensers,
Piping systems. The proportional cost of the piping system in mega plans can be as high as
40% of total investment cost. Incremental cost elements include design, fabrication, associated
valving, installation, insulation, testing, pre-commissioning and initial start-up. A major US pipe
fabrication company is highly aware of these project elements and has promoted induction
heating/bending fabrication techniques to minimize piping bending components and related
Further mega unit features. Operating companies continue to extend capacities on many
similar plants. However, we must recognize that the cost:capacity advantage is less pronounced
at super high capacities, and various associated feedstock, product, utility, emission aspects
pose more difficult operational challenges. Nonetheless, we anticipate continuing plant
capacities in the 1-1.5 MMtpy range where key design elements have been successfully
achieved. Fig. 7 indicates a recent related mega-capacity unit.
Fig. 7 This recently completed mega-unit was the first joint venture ethylene
project and the first unique EPC LSTK unit in China, currently under
A further similar mega-plant petrochemical-refining joint design philosophy has recently been
implemented. Several projects are now under design and construction for Middle East and
South American units based on the concept of enhanced fluid catalytic cracking (FCC) schemes
linked with conventional ethane and /or naphtha type steam crackers. Later in 2008, a mega
plant will start up producing approximately 1.5 MMtpy of ethylene and 1 MMtpy of propylene.
Various product recovery facilities and energy supply/demand systems are integrated, and both
the FCC unit converter system and steam-cracker equipment items incorporate mega-plant
design philosophies. In the role as technology/engineering/construction company, we work
closely with operating companies and key equipment manufacturing companies in order to
implement effective mega-plant cost/capacity designs of joint benefit to all parties—Bigger is
Colin P. Bowen is the vice president of olefins business development for Shaw's
Energy & Chemicals Group (formerly Stone & Webster), where he has worked for
approximately 40 years. He is a graduate of Birmingham University and Imperial
College, UK. Mr. Bowen is a senior member of AIChE, IChemE and IMechE. He
was secretary of AIChE's Fuels & Petrochemicals Division and is Shaw's member
on the AIChE Ethylene Producers Committee. He has worked in the global
petrochemical industry in a range of contractor positions and has contributed
significantly to many ethylene projects during his career.
Daniel Jones is the executive director of furnace technology Shaw's Energy &
Chemicals Group (formerly Stone & Webster) in Houston, Texas. Mr. Jones has
worked in the petroleum industry for 28 years. His varied experience includes
engineered equipment systems for subsea drilling, oil and gas production, gas
pipelines, refining and petrochemical applications. His responsibilities have
included engineering, marketing, sales and executive management. He is a
graduate of the University of Pennsylvania with a BS degree in mechanical