M. de Pontes, R.L. Espinoza, C.P. Nicolaides, J.H. Scholz and M.S. Scurrell (Editors) Natural Gas Conversion IV Studies in Surface Science and Catalysis, Vol. 107 �9 1997 Elsevier Science B.V. All rights reserved.

87

Economic Route for Natural Gas Convers ion to Ethylene and Propy lene

B.V.Vora, a T. L. Marker," P.T. Barger, ' H. R. Nilsen, b S. Kvisle, b and T. Fuglerud b

a UOP, 25 East Algonquin Road, Des Plaines, Illinois, 60017, U.S.A. b Norsk Hydro a.s. Bydoy all6 2, N-0203, Oslo, Norway

1. I N T R O D U C T I O N

The world supply of natural gas continues to increase as a result of the discovery of new fields and the increasing conservation methods being employed in gas and oil recovery. In 1992, the proven gas reserve of the world stood at 145 trillion (1012) cubic meters (tcm), and annual production was 2.5 tcm, an increase of more than 60% from the 1973 production of about 1.6 tcm. During the same period, the percent of gas flared decreased from 13 to 4.4%. However, because the production rate increased, the net quantity of the gas flared was cut only in half, from 0.209 tcm in 1973 to 0.110 tcm in 1992 (Table 1) [ 1 ].

Table 1 World Natural Gas Production and Flaring

Year

1973

1981

1992

Total Production, billions of m 3

1,600

1,800

2,500

Quantity Flared, billions of m 3

209

164

110

Flaring, % of Production

13

The reduction in the quantity flared has been achieved by reinjecting it. Because this process requires compression, it is expensive. Therefore, in many parts of the world where natural gas is in abundance, it is available at a cost as low as $0.30 per million British thermal units (Btu).

The composition of natural gas varies widely, but it consists predominantly of methane. Table 2 shows the composition of a lean and rich natural gas stream. Lean natural gas typically consists of nearly 95% or more methane. The rich natural gas has a significant (15+%) quantity of ethane and heavier fractions. In this case, separating the heavier fractions is economical because liquefied naturals gas (LNG), liquefied petroleum gas (LPG), and condensate have

88

higher values. Methane is most difficult to convert to other higher-value-added products and thus only commands fuel value, which is low where oil and gas reserves are large.

Table 2 Typical Composition of Natural Gas

Ethane Lean, Ethane Rich, Component

N~

CO2

crt,

C2~

C3H8

Ca +

mol-%

0.5

0.2

94.0

4.0

1.0

0.3

mol-%

0.5

1.0

75.0

12.0

8.0

3.5

The efficient conversion of methane to higher-valued products has been a goal of catalytic scientists for the last 50 years. Their efforts divide into two categories: indirect conversion, generally via synthesis gas, and direct conversion. Methanol synthesis and ammonia synthesis, Fischer-Tropsch processes, and Mobil's MTG and MOGD technologies [2] are examples of indirect conversion. The first two processes, which produce high-valued petrochemicals, are widely used in the chemical industry. However, Fisher-Tropsch and the MTG and MOGD technologies, which make liquid transportation fuels, have marginal economics in today's markets [3]. In the last 10 years, major research efforts have also been directed toward the direct conversion of methane. Selective partial oxidation to methanol or other oxygenates [4] and oxidative coupling to higher paraffins and olefins [5] have been the most widely investigated. At the present time, the best yields demonstrated with these technologies are less than 25%, which is below what is required for economic viability [6-10].

Therefore, except for methane conversion to synthesis gas and its subsequent products, such as methanol or ammonia, other developments have not been economically viable under normal market conditions. The challenge is to find an economical means of methane conversion to high-value-added products. Figure 1 shows the relative value that can be obtained from a theoretical conversion of one cubic meter of natural gas (95% methane ) to various products.

Methanol is an important industrial intermediate in the manufacture of. a number of products, such as formaldehyde, acetic acid, and methyl tertiary butyl ether (MTBE). Methanol can also be used as a transportation fuel, but its use has been limited because of its other properties, such as its high water solubility and its blending vapor pressure. Conversion of methanol to gasoline or distillate has been reported [2]. However, as shown in Figure 1, per cubic meter of gas used, these products have lower value than methanol itself. The conversion of methanol to olefins is easier than direct methane coupling and other routes being considered for methane upgrading. Because methanol, though of significant importance, has limited end uses, a large additional production cannot be sustained. The other higher-value products shown in Figure 1 are light olefins (ethylene, propylene).

89

25

20

15

10

Product Value, cents/m3 gas utilized

Natural

Gas

Gasoline

LNG Power

~ o n i ~

Methanol

Uses for Gas

OLEFINS

Figure 1. Value of Products Produced from Natural Gas

2. GAS TO OLEFINS

This paper describes a selective and economical route for converting natural gas to olefins: the GTO process (Figure 2). The first step in the process is natural gas conversion to methanol followed by the UOP/Hydro MTO (methanol to olefins) process using UOP's unique SAPO-34 catalyst. The primary products are ethylene and propylene, two large-volume, highly valued petrochemicals. Table 3 shows 1995 demands for ethylene and propylene and projections for the year 2005. The forecast growth rate amounts to more than 4% per year [6].

Methanol Synthesis Methanol to Olef'ms

Remote ]1 Synthesis ~ SYnthesisGas t MTO Gas "[ Gas to Methanol

,. Ethylene r

Propylene

Figure 2. GTO Process Scheme

90

Table 3 Ethylene and Propylene Demand

Ethylene, Propylene, Year

1995

2000

2005

MM MTA

70

89*

100 ~

MM MTA

37

48*

55*

*Estimates

3. METHANOL SYNTHESIS

The first step of natural gas conversion to methanol is a well-established process. The most widely used processes are:

�9 ICI low-pressure methanol process �9 Lurgi two-step reforming �9 Haldor Topsoe two-step reforming process Recently, Haldor Topsoe has announced a new methanol process development leading to a

single-train capacity of up to 10,000 metric tons per day (MTD) of methanol [12]. To produce 700,000 metric tons per annum (MTA) of ethylene, an MTO unit will need nearly 10,000 metric tons per day (MTD) of methanol feed. The development of a large scale methanol production process makes GTO more attractive.

4. UOP/HYDRO MTO PROCESS

4.1. History of development During the 1980s, scientists at the Molecular Sieve Division of Union Carbide Corporation

(UCC) discovered new molecular sieves, silicoaluminophosphate (SAPO) materials. One of these materials, SAPO-34, was also found to be catalytically very selective for methanol conversion to ethylene and propylene [13,14]. In 1988, Catalysts, Adsorbents and Process Systems of UCC, which included the molecular sieves groups, and the Process Division of UOP Inc. of AlliedSignal merged and formed a partnership company, called UOP, which is jointly owned by AlliedSignal and UCC. UOP continued further development of the SAPO materials as well as their catalytic applications.

During the late 1980s, Norsk Hydro, assisted by Sintef, started independent work on the MTO process. Through this work, Hydro came to know of UOP's SAPO development. For Norway, this process is of a significant importance because Norway has a large gas and oil reserve and a total ban on the flaring of gases from associated oil recovery. In 1992, UOP and Hydro signed an agreement for the joint development of UOP/Hydro MTO process to accelerate commercialization.

The UOP/Hydro MTO process has the advantage of high selectivity to ethylene because of the SAPO-34 catalyst. Ethylene has been shown to be the primary product of this reaction. Ethylene subsequently oligomerizes to higher compounds [ 15, 16], which are thermodynamically

91

favored but are produced at substantially lower levels with the SAPO-34 catalysts than with the ZSM-5 catalysts. Although these compounds may exist within the SAPO-34 cavities, only small linear olefins and paraffins readily pass through the <4 A, diameter pores of the molecular sieve. This high selectivity to ethylene gives SAPO-34 a significant advantage over other types of catalyst systems.

The conversion of methanol to light olefins is done in highly exothermic reactions: 2 CH3OH--> C2H 4 + 2 H20 (AH =-2.8 kcal/mol C at 427~ 3 CH3OH--> C3H6 + 3 H20 (AH =-7.4 kcal/mol C at 427~

The need to remove the high exothermic heat of reaction as well as the need for frequent regeneration led to a fluidized-bed reactor and regenerator design. This design also required development of a fluidized-bed MTO catalyst with the strength and integrity to last in fluidized- bed conditions. UOP has commercially manufactured the MTO catalyst that has shown the type of attrition resistance and stability that enables it to handle multiple regenerations and fluidized- bed conditions.

Multiple regeneration studies in a one-inch fluid-bed reactor were conducted. Figure 3 shows stable activity and selectivity performance after more than 350 regeneration cycles. Small-scale fluidized-bed process variable studies have also been conducted to understand and optimize the MTO catalyst and process.

C2=/C3= 2

.5 -

0.5

. . . . . . . . . . . . . . . . . . . . . . . .

0 100 200

Cycle Number Figure 3. Long-Term Stability

300 400

In addition, a large reactor-regeneration demonstration unit processing methanol feed of 0.5 MTD was built at Norsk Hydro. During 1994, this larger fluid-bed reactor was operated in a batch mode with continuous catalyst withdrawal and addition. The catalyst used was commercially manufactured. Figure 4 shows 50 hours of continuous operation for the large pilot plant. These results confirm the stable operation originally seen in the smaller pilot plant..

At the end of 1994, the demonstration unit was modified to a continuous reaction- regeneration operation by adding a new regeneration vessel and associated catalyst transfer

92

Figure 4. Continuous Test Results

equipment and instrumentation. This unit came on-stream in June 1995 and has been in operation since then. Data from this demonstration unit will be available at the presentation of this paper in November 1995.

Table 4 shows a material balance for the production of 500,000 MTA of ethylene. This amount of ethylene production requires 7,000 MTD of methanol feed. The selectivity of ethylene versus propylene can be changed by adjusting operating conditions. Within reasonable design parameters, one can build flexibility to vary ethylene to propylene ratio from 1:1 to 2:1. Higher ethylene yield increases coke from the yields presented in the base case shown in Table 4.

4.2. Process Flow In the overall flow diagram (Figure 5), the methanol feed is introduced to the fluid-bed

reactor, which is operated to achieve nearly 100% conversion of methanol. The UOP/Hydro MTO process has a fluidized-bed reactor coupled to a fluidized-bed

regenerator. This reactor-regenerator design allows good temperature control and frequent catalyst regeneration. The heat of reaction is removed by steam generation. UOP has extensive commercial experience with fluid-bed catalyst coolers used in the oil refinery fluid catalytic cracking processes [17]. An alternative design involves the installation of cooling coils in the fluidized-bed reactor. Such coils are typically used in the acrylonitrile process or in the production of vinylchloride monomer (VCM).

The spent catalyst is circulated to the regenerator, where coke is burned off, and then returned to the reactor to achieve a steady state. Steam is also generated in the regenerator to remove the exothermic heat from coke burning.

After heat recovery, the reactor effluent is cooled to condense the water produced from the reaction as well as any water present in the crude methanol feed itself.

After water separation, the product gases pass through a caustic scrubber to remove CO2 and then to a dryer. The dry, CO2-free gases are then compressed and processed in a downstream fractionation train. This fractionation section is much simpler than that in a typical steam cracker ethylene unit. No diolefins or acetylene compounds are in the MTO product.

Table 4 MTO Material Balance

Methanol

Ethylene

Propylene

Butylenes

Cs+

H2, C1, C2, C3 Saturates

COx

Coke

Water

Total

Feed, MTA

2,330

2,330

Products, MTA

500

327

100

22

35

31

1,310

2,330

% Yield on C

49

32

10

3.5

0.5

3.0

100

93

Figure 5. MTO Process

94

Because the product has little hydrogen or methane, demethanization is much simpler. The concentration of ethane and propane is also low. In fact, chemical-grade propylene can be produced without a propane-propylene splitter (Table 4).

Because the MTO process for the production of ethylene and propylene requires a large quantity of methanol feed, the methanol production will be at the same site as the MTO process in many cases. Significant possibilities exist for process integration with methanol production and purification. This integration can result in substantial savings (20%) in investment.

4.3. Economics Three cases for the production of 500,000 MTA of ethylene, shown in Table 5, were

developed for purposes of economic comparison: �9 Naphtha cracker �9 Ethane cracker �9 Gas to olefins (GTO) process

Table 5 Material Balances for Na ththa Cracking, Ethane Cracking and GTO Routes

10 3 MTA

Naphtha Ethane Feed and Products Cracking Cracking GTO �9

Feed:

Naphtha Ethane Natural Gas Water TOTAL

1,380

1,380

620 - -

- -

620

1,300 1,700 3,000

Products"

Ethylene Propylene C4's C5+ HDT Pygas Gas HE, CI-C4, Fuel Gas, Coke Fuel Oil Water TOTAL

5OO 203 138

252 230 57

1380

5OO 10 2O

9O

620

500 325 100 25 _ _

35O

1,700 3,000

95

The information on the naphtha and ethane crackers is based on published data [ 11,18]. The feedstock requirement for producing 500,000 MTA of ethylene per these routes are:

�9 Naphtha: 1.38 MM MTA �9 Ethane: 0.62 MM MTA (4.1 MM MTA rich natural gas) �9 Natural gas for GTO: 1.30 MM MTA Table 6 shows the basis for the economic comparison. For all the cases, the natural gas

pricing and capital investment are for a remote site location. Table 7 compares investment requirements on the U.S. Gulf Coast versus a remote site. Table 8 compares the three cases.

Table 6 GTO Economics Basis Remote Site

Feed and Product i Value

Feed"

Naphtha

Ethane

Natural Gas

Products:

Ethylene

Propylene

C4's

H2+ Fuel Gas

Hdt. Pygas

Fuel Oil

$170/MT ($20.00/bbl)

$45/MT ($1.00/MM Btu)

$24/MT ($0.50/MM Btu)

$500/MT ($0.23/1b)

$400/MT ($0.18/lb)

$170/MT ($0.37/gal)

$24/MT ($0.50/MM Btu)

$200/MT ($23.0/bbl)

$100/MT ($0.27/gal)

Table 7 Capital Investment Costs

Gulf Coast, 2 nd Quarter 1995

Remote Site, 2 nd Quarter 1995

Naphtha Cracking

650

900

$ M M

Ethane Cracking

450

600

GTO

830

1,100

96

As shown in Table 8 and Figure 6, the total variable cost (feed minus by-products plus utilities, catalyst, and chemicals) for the GTO process is lower by $371/MT relative to a naphtha cracker and is lower by $251/MT relative to an ethane cracker. The main by-product of the GTO process is propylene. These economics used the historical value of $400/MT for propylene and not the current high market price of $600/MT prevailing throughout the world in 1995 [19].

As shown in Table 8, alter consideration of fixed charges and return on investment, the ethane cracking appears to be more attractive than the GTO process. Naphtha cracking is the least-attractive option. However, getting large quantities of ethane at such low feedstock value is not realistic. Ethane cracking requires investment in an NGL plant, which has not been included in the capital cost. In addition to the recovery of ethane, heavier fractions must be recovered and marketed in the ethane cracking case. To produce 620,000 MTA of ethane, 4.1 MM MTA of a rich natural gas stream (containing 15% ethane) would have to be processed. More than 3 MM MTA of lean methane gas would then be available for other uses, which are quite limited. The result is significant additional cost for reinjection, as the flaring option is environmentally unacceptable. If these associated costs of ethane recovery are considered, the economics of the GTO route becomes the most attractive.

Table 8 Economic Summary

Capital Investment 1st Qtr. 1995, $MM, U.S., Remote Site

Production Cost, $/MT C2-: Feed By-product Credit Net Raw Materials

Utilities, Catalyst, & Chemicals Total Variable Cost

Fixed Cost @ 10% of Inv. Total Cost of Production

Sale Price Margin

Simple ROI before Taxes

Naptha Cracking

900

470 -333 137

38 175

180 355

500 145

Ethane Cracking

600 ~

56 -13 43

12 55

120 175

500 325

27

GTO

1,100

67 -318 -251

55 -196

220 24

500 476

22

* Does not include investment of LNG plant for ethane recovery for natural gas.

97

Variable costs, $/MT C2 = 200

100

-100

-200

-300

Naphtha Cracking Ethane Cracking

GTO

Processing Routes

Figure 6. Variable Cost of Ethylene Production

4.4. Increased Production from Naphtha Crackers

As discussed earlier and as shown in Table 5, the MTO reaction after the separation of water produces relatively high-purity ethylene, propylene, and butylene. Because this stream has little saturates, hydrogen, and methane, it is easy to add to the inlet of the existing compressor of a naphtha cracker. With minor modification of the downstream distillation section of a naphtha cracker, an additional 20 to 30% ethylene and propylene capacity can be achieved. The MTO reaction and regenerator are easy to scale up or down for a wide range of production.

4.5. Small-Scale Ethylene Production

Some developing nations require small ethylene-producing sites to meet local demand. However, in many places, the transportation of ethylene is expensive and so is not feasible. Methanol can be shipped to these sites, and MTO units can be built to meet the local ethylene demand.

5. CONCLUSION

The UOP/Hydro MTO process and catalyst have been successfully developed and are currently available for license from UOP and Hydro. This technology allows the production of high-valued ethylene from natural gas. The GTO UOP/Hydro MTO process has favorable economics in areas where low-cost natural gas is available and easily beats the internal rates of return from traditional naphtha cracking in these locations.

98

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