Chapter 5.1 - UOP Oleflex Process for Light Olefin Production

UOP OLEFLEX PROCESS FOR LIGHT OLEFIN PRODUCTION

5.4



DEHYDROGENATION



FIGURE 5.1.1



Oleflex process flow.



The reactor section consists of several radial-flow reactors, charge and interstage

heaters, and a reactor feed-effluent heat exchanger. The diagram shows a unit with four

reactors, which would be typical for a unit processing propane feed. Three reactors are

used for butane or isopentane dehydrogenation. Three reactors are also used for blends of

C3-C4 or C4-C5 feeds.

Because the reaction is endothermic, conversion is maintained by supplying heat

through interstage heaters. The effluent leaves the last reactor, exchanges heat with the

combined feed, and is sent to the product recovery section.



Product Recovery Section

A simplified product recovery section is also shown in Fig. 5.1.1. The reactor effluent

is cooled, compressed, dried, and sent to a cryogenic separation system. The dryers

serve two functions: (1) to remove trace amounts of water formed from the catalyst

regeneration and (2) to remove hydrogen sulfide. The treated effluent is partially condensed in the cold separation system and directed to a separator.

Two products come from the Oleflex product recovery section: separator gas and separator liquid. The gas from the cold high-pressure separator is expanded and divided into

two streams: recycle gas and net gas. The net gas is recovered at 90 to 93 mol % hydrogen

purity. The impurities in the hydrogen product consist primarily of methane and ethane.

The separator liquid, which consists primarily of the olefin product and unconverted paraffin, is sent downstream for processing.



Catalyst Regeneration Section

The regeneration section, shown in Fig. 5.1.2, is similar to the CCR* unit used in the

UOP Platforming* process. The CCR unit performs four functions:

●

●



Burns the coke off the catalyst

Redistributes the platinum

*Trademark and/or service mark of UOP.



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UOP OLEFLEX PROCESS FOR LIGHT OLEFIN PRODUCTION

UOP OLEFLEX PROCESS FOR LIGHT OLEFINS



FIGURE 5.1.2



●

●



5.5



Oleflex regeneration section.



Removes the excess moisture

Reduces the catalyst prior to returning to the reactors



The slowly moving bed of catalyst circulates in a loop through the reactors and the

regenerator. The cycle time around the loop can be adjusted within broad limits but is typically anywhere from 5 to 10 days, depending on the severity of the Oleflex operation and

the need for regeneration. The regeneration section can be stored for a time without interrupting the catalytic dehydrogenation process in the reactor and recovery sections.



DEHYDROGENATION PLANTS

Propylene Plant

Oleflex process units typically operate in conjunction with fractionators and other

process units within a production plant. In a propylene plant (Figure 5.1.3), a propanerich liquefied petroleum gas (LPG) feedstock is sent to a depropanizer to reject butanes

and heavier hydrocarbons. The depropanizer overhead is then directed to the Oleflex

unit. The once-through conversion of propane is approximately 40 percent, which

closely approaches the equilibrium value defined by the Oleflex process conditions.

Approximately 90 percent of the propane conversion reactions are selective to propylene and hydrogen; the result is a propylene mass selectivity in excess of 85 wt %. Two

product streams are created within the C3 Oleflex unit: a hydrogen-rich vapor product

and a liquid product rich in propane and propylene.

Trace levels of methyl acetylene and propadiene are removed from the Oleflex liquid

product by selective hydrogenation. The selective diolefin and acetylene hydrogenation

step is accomplished with the Hüls SHP process, which is available for license through

UOP. The SHP process selectively saturates diolefins and acetylenes to monoolefins



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UOP OLEFLEX PROCESS FOR LIGHT OLEFIN PRODUCTION

5.6



DEHYDROGENATION



FIGURE 5.1.3



C3 Oleflex plant.



without saturating propylene. The process consists of a single liquid-phase reactor. The

diolefins plus acetylene content of the propylene product is less than 5 wt ppm.

Ethane and lighter material enter the propylene plant in the fresh feed and are also created by nonselective reactions within the Oleflex unit. These light ends are rejected from

the complex by a deethanizer column. The deethanizer bottoms are then directed to a

propane-propylene (P-P) splitter. The splitter produces high-purity propylene as the overhead product. Typical propylene purity ranges between 99.5 and 99.8 wt %. Unconverted

propane from the Oleflex unit concentrates in the splitter bottoms and is returned to the

depropanizer for recycle to the Oleflex unit.



Ether Complex

A typical etherification complex configuration is shown in Fig. 5.1.4 for the production

of methyl tertiary butyl ether (MTBE) from butanes and methanol. Ethanol can be substituted for methanol to make ethyl tertiary butyl ether (ETBE) with the same process

configuration. Furthermore, isopentane may be used in addition to or instead of field

butanes to make tertiary amyl methyl ether (TAME) or tertiary amyl ethyl ether

(TAEE). The complex configuration for a C5 dehydrogenation complex varies according to the feedstock composition and processing objectives.

Three primary catalytic processes are used in an MTBE complex:

●

●

●



Paraffin isomerization to convert normal butane into isobutane

Dehydrogenation to convert isobutane into isobutylene

Etherification to react isobutylene with methanol to make MTBE



Field butanes, a mixture of normal butane and isobutane obtained from natural gas condensate, are fed to a deisobutanizer (DIB) column. The DIB column prepares an isobutane

overhead product, rejects any pentane or heavier material in the DIB bottoms, and makes

a normal butane sidecut for feed to the paraffin isomerization unit.

The DIB overhead is directed to the Oleflex unit. The once-through conversion of

isobutane is approximately 50 percent. About 91 percent of the isobutane conversion reactions are selective to isobutylene and hydrogen. On a mass basis, the isobutylene selectivity



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UOP OLEFLEX PROCESS FOR LIGHT OLEFIN PRODUCTION

UOP OLEFLEX PROCESS FOR LIGHT OLEFINS



FIGURE 5.1.4



5.7



MTBE production facility.



is 88 wt %. Two product streams are created within the C4 Oleflex unit: a hydrogen-rich

vapor product and a liquid product rich in isobutane and isobutylene.

The C4 Oleflex liquid product is sent to an etherification unit, where methanol reacts

with isobutylene to make MTBE. Isobutylene conversion is greater than 99 percent, and

the MTBE selectivity is greater than 99.5 percent. Raffinate from the etherification unit is

depropanized to remove propane and lighter material. The depropanizer bottoms are then

dried, saturated, and returned to the DIB column.



PROPYLENE PRODUCTION ECONOMICS

A plant producing 350,000 MTA of propylene is chosen to illustrate process economics. Given the more favorable C4 and C5 olefin equilibrium, butylene and amylene production costs are lower per unit of olefin when adjusted for any differential in feedstock

value. The basis used for economic calculations is shown in Table 5.1.1. This basis is

typical for U.S. Gulf Coast prices prevailing in mid-2002 and can be used to show that

the pretax return on investment for such a plant is approximately 24 percent.



Material Balance

The LPG feedstock is the largest cost component of propylene production. The quantity of

propane consumed per unit of propylene product is primarily determined by the selectivity of the Oleflex unit because fractionation losses throughout the propylene plant are small.

The Oleflex selectivity to propylene is 90 mol % (85 wt %), and the production of 1.0 metric ton (MT) of propylene requires approximately 1.2 MT of propane.

An overall mass balance for the production of polymer-grade propylene from C3 LPG

is shown in Table 5.1.2 for a polymer-grade propylene plant producing 350,000 MTA,

based on 8000 operating hours per year. The fresh LPG feedstock is assumed to be 94 LV %

propane with 3 LV % ethane and 3 LV % butane. The native ethane in the feed is rejected in

the deethanizer along with light ends produced in the Oleflex unit and used as process fuel.

The butanes are rejected from the depropanizer bottoms. This small butane-rich



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UOP OLEFLEX PROCESS FOR LIGHT OLEFIN PRODUCTION

5.8



DEHYDROGENATION



TABLE 5.1.1

Calculations



Utility, Feed, and Product Valuations for Economic



Utility values

Fuel gas

Boiler feed water

Cooling water

Electric power



$2.80/million Btu

$0.45/klb

$0.12/kgal

$0.05/kWh



$11.10/million kcal

$1.00/MT

$0.03/m3

$0.05/kWh



Feed and product values

C3 LPG (94 LU % propane)

Propylene (99.5 wt %)

Note:



$0.35/gal

$0.19/lb



$180/MT

$420/MT



MT ϭ metric tons; SCF ϭ standard cubic feet.



TABLE 5.1.2

Plant



Material Balance for a 350,000-MTA Propylene

Flow rate,

MT/h



Flow rate,

MTA



Feed:

C3 LPG (94 LV % propane)



55.00



440,000



Products:

Propylene (99.5 wt %)

Fuel by-products



43.75

11.25



350,000

90,000



55.00



440,000



Total products

Note:



MT/h ϭ metric tons per hour; MTA ϭ metric tons per annum.



stream could be used as either a by-product or as fuel. In this example, the depropanizer

bottoms were used as fuel within the plant.

The Oleflex process coproduces high-quality hydrogen. Project economics benefit

when a hydrogen consumer is available in the vicinity of the propylene plant. If chemical

hydrogen cannot be exported, then hydrogen is used as process fuel. This evaluation

assumes that hydrogen is used as fuel within the plant.



Utility Requirements

Utility requirements for a plant producing 350,000 MTA of propylene are summarized

in Table 5.1.3. These estimates are based on the use of an extracting steam turbine to

drive the Oleflex reactor effluent compressor. A water-cooled surface condenser is used

on the steam turbine exhaust. A condensing steam driver was chosen in this example for

the propane-propylene splitter heat-pump compressor.



Propylene Production Costs

Representative costs for producing 350,000 MTA of polymer-grade propylene using the

Oleflex process are shown in Table 5.1.4. These costs are based on feed and product



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UOP OLEFLEX PROCESS FOR LIGHT OLEFIN PRODUCTION

5.9



UOP OLEFLEX PROCESS FOR LIGHT OLEFINS



TABLE 5.1.3 Net Utility Requirements for a 350,000-MTA

Propylene Plant

Utility cost

Utility requirements

Electric power

Boiler feed water

Cooling water

Fuel gas

Net utilities

Note:



Consumption



$/h



$/MTA C3



6,500 kW

10 MT/h

6,000 m3/h

(13.1 million kcal/h)



325

10

180

145



7.43

0.23

4.11

3.31

15.08



MTA ϭ metric tons per annum; MT/h ϭ metric tons per hour.



TABLE 5.1.4 Cost for Producing 350,000 MTA of Polymer-Grade

Propylene Using the Oleflex Process

Costs

Revenues,

million $/year

Propylene product

Propane feedstock

Net utilities

Catalyst and chemicals

Fixed expenses

Total

Note:



million

$/year



$/MT C3



147.0

—

—

—

—

147.0



—

79.2

5.3

3.8

7.0

95.3



—

226.3

15.1

10.9

20.0

272.3



MTA ϭ metric tons per annum; MT ϭ metric tons.



values defined in Table 5.1.1. The fixed expenses in Table 5.1.4 consist of estimated

labor costs and maintenance costs and include an allowance for local taxes, insurance,

and interest on working capital.



Capital Requirements

The ISBL erected cost for an Oleflex unit producing 350,000 MTA of polymer-grade

propylene is approximately $145 million (U.S. Gulf Coast, mid-2002 erected cost).

This figure includes the reactor and product recovery sections, a modular CCR unit, a

Hüls SHP unit, and a fractionation section consisting of a depropanizer, deethanizer,

and heat-pumped P-P splitter. The costs are based on an extracting steam turbine driver for the reactor effluent compressor and a steam-driven heat pump. Capital costs are

highly dependent on many factors, such as location, cost of labor, and the relative workload of equipment suppliers.

Total project costs include ISBL and OSBL erected costs and all owner’s costs. This

example assumes an inclusive mid-2002 total project cost of $215 million including:

●

●

●



ISBL erected costs for all process units

OSBL erected costs (off-site utilities, tankage, laboratory, warehouse, for example)

Initial catalyst and absorbant loadings



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UOP OLEFLEX PROCESS FOR LIGHT OLEFIN PRODUCTION

5.10

●

●



DEHYDROGENATION



Technology fees

Project development including site procurement and preparation



Overall Economics

Because the feedstock represents such a large portion of the total production cost, the

economics for the Oleflex process are largely dependent on the price differential

between propane and propylene. Assuming the values of $180/MT for propane and

$420/MT for propylene, or a differential price of $240/MT, the pretax return on investment is approximately 24 percent for a plant producing 350,000 MTA of propylene.



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Source: HANDBOOK OF PETROLEUM REFINING PROCESSES



CHAPTER 5.2



UOP PACOL

DEHYDROGENATION

PROCESS

Peter R. Pujadó

UOP LLC

Des Plaines, Illinois



INTRODUCTION

Paraffins can be selectively dehydrogenated to the corresponding monoolefins by using

suitable dehydrogenation catalysts. Iron catalysts have long been used for the dehydrogenation of ethylbenzene to styrene, and catalysts made of chromia (chrome oxide) supported on alumina have long been used for the dehydrogenation of light paraffins (for

example, n-butane to n-butene) and the deeper dehydrogenation of olefins to diolefins

(for example, n-butene to 1,3-butadiene). However, newer commercial processes for the

dehydrogenation of light and heavy paraffins are based on the use of noble-metal catalysts because of the superior stability and selectivity of these catalyst systems.

In the late 1940s and through the 1950s, the pioneering work done at UOP* by Vladimir

Haensel on platinum catalysis for the catalytic reforming of naphthas for the production of

high-octane gasolines and high-purity aromatics showed that platinum catalysts have interesting dehydrogenation functions. This research area was later pursued by Herman Bloch and

others also within UOP. In 1963-64, UOP started development work on heterogeneous platinum catalysts supported on an alumina base for the dehydrogenation of heavy n-paraffins.

The resulting successful process, known as the Pacol* process (for paraffin conversion to

olefins), was first commercialized in 1968. The advent of the UOP Pacol process marked a

substantial transformation in the detergent industry and contributed to the widespread use of

linear alkylbenzene sulfonate (LAS or LABS) on an economical, cost-effective basis. As of

mid-2003, more than 40 Pacol units have been built, or are under design or construction;

practically all new linear alkylbenzene (LAB) capacity built on a worldwide basis over the

last two decades makes use of UOP’s Pacol catalytic dehydrogenation process.

Maintaining technological superiority over some 30-odd years requires continued innovation and improvement, principally of the dehydrogenation catalyst, the reactor design,

and operating conditions because these have the greatest impact on the overall process economics. The first commercial Pacol dehydrogenation catalysts, denoted DeH-3 and DeH*Trademark and/or service mark of UOP.



5.11

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UOP PACOL DEHYDROGENATION PROCESS

5.12



DEHYDROGENATION



4, came on-stream in the mid-1960s. They were soon superseded by a newer catalyst, DeH5, that was commercialized in 1971 and dominated the market for several years. In 1983,

DeH-7 catalyst was introduced. This new catalyst exhibited about 1.75 times the stability

of its predecessor, DeH-5, and soon replaced it as the dominant Pacol catalyst.

Development efforts continued, and in 1998, DeH-11 was commercialized. This catalyst is

the first “layered sphere” catalyst to be offered by UOP in which a thin reactive layer is

coated onto an inert one. The result is an advantage in selectivity to mono-olefins. In 2001

DeH-201 was introduced. This catalyst, also a layered sphere, allows for higher conversion

operation than previous Pacol catalysts. All these various generations of paraffin dehydrogenation catalysts have resulted in improved yields at higher conversion and higher operating severities, thus allowing for smaller and more economical units for a given

production capacity.

Since 1980, UOP has adapted similar catalysts to the selective catalytic dehydrogenation of light olefins (propane to propylene and isobutane to isobutylene) in the Oleflex*

process; a number of large-capacity units have been built for this application. Because of

the higher severity, light paraffin dehydrogenation units make use of UOP’s proprietary

CCR* continuous catalyst regeneration technology, which was originally developed and

commercialized for the catalytic reforming of naphthas at high severity. Because the Pacol

process operates at a lower severity, catalyst runs are significantly longer, and CCR technology is not needed.



PROCESS DESCRIPTION

The catalytic reaction pathways found in the dehydrogenation of n-paraffins to nmonoolefins [linear internal olefins (LIO)] in addition to other thermal cracking reactions are illustrated in Fig. 5.2.1. A selective catalyst is required if only LIO is to be the

main product.

In the Pacol reaction mechanism, the conversion of n-paraffins to monoolefins is near

equilibrium, and therefore a small but significant amount of diolefins and aromatics is produced. In the alkylation process, the diolefins consume 2 moles of benzene to yield heavier diphenylalkane compounds or form heavier polymers that become part of the heavy

alkylate and the bottoms by-products of the hydrofluoric (HF) acid regenerator. Thus,



FIGURE 5.2.1



Dehydrogenation reaction pathways.



*Trademark and/or service mark of UOP.



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