Cascade refrigeration process for liquefaction of natural gas

This invention relates to a process for liquefying a pressurized gas stream rich in methane in which the liquefication of the gas stream occurs in a heat exchanger being cooled by a cascade refrigeration system to produce a methane-rich liquid product having a temperature above about -112.degree. C. (-170.degree. F.). In this process, a pressurized gas stream is introduced into heat exchange contact with a first refrigerant cycle comprising at least one refrigeration stage whereby the temperature of the gas stream is reduced by heat exchange with a first portion of a first refrigerant to produce a cooled gas stream. The cooled gas stream is then introduced into heat exchange contact with a second refrigerant cycle comprising at least one refrigeration stage whereby the temperature of the cooled gas stream is further reduced by heat exchange with a second refrigerant to produce a liquefied methane-rich stream having a temperature above about -112.degree. C. (-170.degree. F.) and a pressure sufficient for the liquefied stream to be at or below its bubble point.

FIELD OF THE INVENTION 
This invention relates to a natural gas liquefaction process, and more 
particularly relates to a process to produce pressurized liquid natural 
gas (PLNG). 
BACKGROUND OF THE INVENTION 
Because of its clean burning qualities and convenience, natural gas has 
become widely used in recent years. Many sources of natural gas are 
located in remote areas, great distances from any commercial markets for 
the gas. Sometimes a pipeline is available for transporting produced 
natural gas to a commercial market. When pipeline transportation is not 
feasible, produced natural gas is often processed into liquefied natural 
gas (which is called "LNG") for transport to market. 
One of the distinguishing features of a LNG plant is the large capital 
investment required for the plant. The equipment used to liquefy natural 
gas is generally quite expensive. The liquefaction plant is made up of 
several basic systems, including gas treatment to remove impurities, 
liquefaction, refrigeration, power facilities, and storage and ship 
loading facilities. While the cost of LNG plant can vary widely depending 
upon plant location, a typical conventional LNG project can cost from U.S. 
$5 billion to U.S. $10 billion, including field development costs. The 
plant's refrigeration systems can account for up to 30 percent of the 
cost. 
In the design of a LNG plant, three of the most important considerations 
are (1) the selection of the liquefaction cycle, (2) the materials used in 
the containers, piping, and other equipment, and (3) the process steps for 
converting a natural gas feed stream into LNG. 
LNG refrigeration systems are expensive because so much refrigeration is 
needed to liquefy natural gas. A typical natural gas stream enters a LNG 
plant at pressures from about 4,830 kPa (700 psia) to about 7,600 kPa 
(1,100 psia) and temperatures from about 20.degree. C. (68.degree. F.) to 
about 40.degree. C. (104.degree. F.). Natural gas, which is predominantly 
methane, cannot be liquefied by simply increasing the pressure, as is the 
case with heavier hydrocarbons used for energy purposes. The critical 
temperature of methane is -82.5.degree. C. (-116.5.degree. F.). This means 
that methane can only be liquefied below that temperature regardless of 
the pressure applied. Since natural gas is a mixture of gases, it 
liquefies over a range of temperatures. The critical temperature of 
natural gas is between about -85.degree. C. (-121.degree. F.) and 
-62.degree. C. (-80.degree. F.). Typically, natural gas compositions at 
atmospheric pressure will liquefy in the temperature range between about 
-165.degree. C. (-265.degree. F.) and -155.degree. C. (-247.degree. F.). 
Since refrigeration equipment represents such a significant part of the 
LNG facility cost, considerable effort has been made to reduce 
refrigeration costs. 
Although many refrigeration cycles have been used to liquefy natural gas, 
the three types most commonly used in LNG plants today are: (1) "expander 
cycle" which expands gas from a high pressure to a low pressure with a 
corresponding reduction in temperature, (2) "multi-component refrigeration 
cycle" which uses a multi-component refrigerant in specially designed 
exchangers, and (3) "cascade cycle" which uses multiple single component 
refrigerants in heat exchangers arranged progressively to reduce the 
temperature of the gas to a liquefaction temperature. Most natural gas 
liquefaction cycles use variations or combinations of these three basic 
types. 
The cascade system generally uses two or more refrigeration loops in which 
the expanded refrigerant from one stage is used to condense the compressed 
refrigerant in the next stage. Each successive stage uses a lighter, more 
volatile refrigerant which, when expanded, provides a lower level of 
refrigeration and is therefore able to cool to a lower temperature. To 
diminish the power required by the compressors, each refrigeration cycle 
is typically divided into several pressure stages (three or four stages is 
common). The pressure stages have the effect of dividing the work of 
refrigeration into several temperature steps. Propane, ethane, ethylene, 
and methane are commonly used refrigerants. Since propane can be condensed 
at a relatively low pressure by air coolers or water coolers, propane is 
normally the first-stage refrigerant. Ethane or ethylene can be used as 
the second-stage refrigerant. Condensing the ethane exiting the ethane 
compressor requires a low-temperature coolant. Propane provides this 
low-temperature coolant function. Similarly, if methane is used as a 
final-stage coolant, ethane is used to condense methane exiting the 
methane compressor. The propane refrigeration system is therefore used to 
cool the feed gas and to condense the ethane refrigerant and ethane is 
used to further cool the feed gas and to condense the methane refrigerant. 
Materials used in conventional LNG plants also contribute to the plants' 
cost. Containers, piping, and other equipment used in LNG plants are 
typically constructed, at least in part, from aluminum, stainless steel or 
high nickel content steel to provide the necessary strength and fracture 
toughness at low temperatures. 
In conventional LNG plants water, carbon dioxide, sulfur-containing 
compounds, such as hydrogen sulfide and other acid gases, n-pentane and 
heavier hydrocarbons, including benzene, must be substantially removed 
from the natural gas processing, down to parts-per-million (ppm) levels. 
Some of these compounds will freeze, causing plugging problems in the 
process equipment. Other compounds, such as those containing sulfur, are 
typically removed to meet sales specifications. In a conventional LNG 
plant, gas-treating equipment is required to remove the carbon dioxide and 
acid gases. The gas treating equipment typically uses a chemical and/or 
physical solvent regenerative process and requires a significant capital 
investment. Also, the operating expenses are high. Dry bed dehydrators, 
such as molecular sieves, are required to remove the water vapor. A scrub 
column and fractionation equipment are typically used to remove the 
hydrocarbons that tend to cause plugging problems. Mercury is also removed 
in a conventional LNG plant since it can cause failures in equipment 
constructed of aluminum. In addition, a large portion of the nitrogen that 
may be present in natural gas is removed after processing since nitrogen 
will not remain in the liquid phase during transport of conventional LNG 
and having nitrogen vapor in LNG containers at the point of delivery is 
undesirable. 
There is a continuing need in the industry for an improved process for 
liquefying natural gas which minimizes the amount of refrigeration 
equipment and the process horsepower required. 
SUMMARY 
This present invention relates generally to a liquefaction process of a gas 
stream rich in methane and having an initial pressure above about 3,100 
kPa (450 psia). The primarily refrigeration for condensing the natural gas 
is by cascade refrigeration cycles, preferably only two cycles. The 
natural gas is then pressure expanded by a suitable pressure expansion 
means to produce a methane-rich liquid product having a temperature above 
about -112.degree. C. (-170.degree. F.) and a pressure sufficient for the 
liquid product to be at or below its bubble point. 
The process of this invention may also condense boil-off vapor produced by 
a pressurized liquid natural gas. If the natural gas contains hydrocarbons 
heavier than methane and it is desired to remove the heavier hydrocarbons, 
a fractionation process may be added to the process. 
The method of the present invention can be used both for the initial 
liquefaction of a natural gas at the source of supply for storage or 
transportation, and to re-liquefy natural gas vapor given off during 
storage and ship loading. Accordingly, an object of this invention is to 
provide an improved liquefaction system for the liquefaction or 
reliquefaction of natural gas. Another object of this invention is to 
provide an improved liquefaction system wherein substantially less 
compression power is required than in prior art systems. A still further 
object of the invention is to provide an improved liquefaction process 
that is economical and efficient in operation. The very low temperature 
refrigeration of conventional LNG process is very expensive compared to 
the relatively mild refrigeration needed in the production of PLNG in 
accordance with the practice of this invention.

DESCRIPTION OF THE PREFERRED EMBODIMENTS 
The present invention uses a cascade refrigeration system to liquefy 
natural gas to produce a methane-rich liquid product having a temperature 
above about -112.degree. C. (-170.degree. F.) and a pressure sufficient 
for the liquid product to be at or below its bubble point. This 
methane-rich product is sometimes referred to in this description as 
pressurized liquid natural gas (PLNG). The term "bubble point" is the 
temperature and pressure at which a liquid begins to convert to gas. For 
example, if a certain volume of PLNG is held at constant pressure, but its 
temperature is increased, the temperature at which bubbles of gas begin to 
form in the PLNG is the bubble point. Similarly, if a certain volume of 
PLNG is held at constant temperature but the pressure is reduced, the 
pressure at which gas begins to form defines the bubble point. At the 
bubble point the mixture is a saturated liquid. 
Using a cascade refrigeration system in accordance with the present 
invention requires less power for liquefying the natural gas than cascade 
refrigeration processes used in the past and the equipment used in the 
process of this invention can be made of less expensive materials. By 
contrast, prior art processes that produce LNG at atmospheric pressures 
having temperatures as low as -160.degree. C. (-256.degree. F.) require 
that at least part of the process equipment be made of expensive materials 
for safe operation. 
The energy needed for liquefying the natural gas in the practice of this 
invention is greatly reduced over energy requirements of a conventional 
LNG plant. The reduction in necessary refrigeration energy required for 
the process of the present invention results in a large reduction in 
capital costs, proportionately lower operating expenses, and increased 
efficiency and reliability, thus greatly enhancing the economics of 
producing liquefied natural gas. 
At the operating pressures and temperatures of the present invention, about 
3 1/2 weight percent nickel steel can be used in piping and facilities in 
the coldest operating areas of the liquefaction process, whereas the more 
expensive 9 weight percent nickel or aluminum is generally required for 
the same equipment in a conventional LNG process. This provides another 
significant cost reduction for the process of this invention compared to 
prior art LNG processes. 
The first consideration in cryogenic processing of natural gas is 
contamination. The raw natural gas feed stock suitable for the process of 
this invention may comprise natural gas obtained from a crude oil well 
(associated gas) or from a gas well (non-associated gas). The composition 
of natural gas can vary significantly. As used herein, a natural gas 
stream contains methane (C.sub.1) as a major component. The natural gas 
will typically also contain ethane (C.sub.2), higher hydrocarbons 
(C.sub.3+), and minor amounts of contaminants such as water, carbon 
dioxide, hydrogen sulfide, nitrogen, butane, hydrocarbons of six or more 
carbon atoms, dirt, iron sulfide, wax, and crude oil. The solubilities of 
these contaminants vary with temperature, pressure, and composition. At 
cryogenic temperatures, CO.sub.2, water, and other contaminants can form 
solids, which can plug flow passages in cryogenic heat exchangers. These 
potential difficulties can be avoided by removing such contaminants if 
conditions within their pure component, solid phase temperature-pressure 
phase boundaries are anticipated. In the following description of the 
invention, it is assumed that the natural gas stream has been suitably 
treated to remove sulfides and carbon dioxide and dried to remove water 
using conventional and well known processes to produce a "sweet, dry" 
natural gas stream. If the natural gas stream contains heavy hydrocarbons 
which could freeze out during liquefaction or if the heavy hydrocarbons 
are not desired in the PLNG, the heavy hydrocarbon may be removed by a 
fractionation process prior to producing the PLNG as described in more 
detail below. 
One advantage of the present invention is that the warmer operating 
temperatures enables the natural gas to have higher concentration levels 
of freezable components than would be possible in a conventional LNG 
process. For example, in a conventional LNG plant that produces LNG at 
-160.degree. C. (-256.degree. F.) the CO.sub.2 must be below about 50 ppm 
to avoid freezing problems. In contrast, by keeping the process 
temperatures above about -112.degree. C. (-170.degree. F.), the natural 
gas can contain CO.sub.2 at levels as high as about 1.4 mole % CO.sub.2 at 
temperatures of -112.degree. C. (-170.degree. F.) and about 4.2% at 
-95.degree. C. (-139.degree. F.) without causing freezing problems in the 
liquefaction process of this invention. 
Additionally, moderate amounts of nitrogen in the natural gas need not be 
removed in the process of this invention because nitrogen will remain in 
the liquid phase with the liquefied hydrocarbons at the operating 
pressures and temperatures of the present invention. The ability to 
reduce, or in some cases omit, the equipment required for gas treating and 
nitrogen rejection when the composition of the natural gas allows, 
provides significant technical and economic advantages. These and other 
advantages of the invention will be better understood by referring to the 
Figures. 
Referring to FIG. 1, pressurized natural gas feed stream 10 preferably 
enters the liquefaction process at a pressure above about 1,724 kPa (250 
psia) and more preferably above about 4,830 kPa (700 psia) and preferably 
at temperatures below about 40.degree. C. (104.degree. F.); however, 
different pressures and temperatures can be used, if desired, and the 
system can be appropriately modified accordingly by persons skilled in the 
art taking into account the teachings of this invention. If the gas stream 
10 is below about 1,724 kPa (250 psia), it can be pressurized by a 
suitable compression means (not shown), which may comprise one or more 
compressors. 
The feed stream 10 passes through a series of heat exchangers, preferably 
two heat exchangers 30 and 31, which are refrigerated by a first 
refrigeration cycle 32. Refrigeration cycle 32 cools the feed stream 10 in 
heat exchangers 30 and 31 and cools refrigerant in a second refrigeration 
cycle 33 which is downstream in the liquefaction process. Refrigeration 
cycle 33 further cools the natural gas in a series of heat exchangers, 
preferably three exchangers 37, 38, and 39 as shown in FIG. 1. The design 
and operation of the refrigeration cycles 32 and 33 are well known to 
those skilled in the art and details of their operation are found in the 
prior art. The refrigerant in the first refrigeration cycle 32 is 
preferably propane and the refrigerant in the second refrigeration cycle 
33 is preferably ethylene. Examples of cascade refrigeration systems are 
described in U.S. Pat. No. 3,596,472; Plant Processing of Natural Gas, 
issued by the Petroleum Extension Service, The University of Texas at 
Austin, Tex. (1974); and Harper, E. A. et. al., Trouble Free LNG, Chemical 
Engineering Progress, Vol. 71, No. 11 (1975). 
Liquefied natural gas stream 19 exiting the last heat exchanger 39 in 
accordance with the practice of this invention has a temperature above 
-112.degree. C. (-170.degree. F.) and a pressure sufficient for the liquid 
product to be at or below its bubble point. If the pressure of stream 10 
as it exits the last stage of the second refrigeration cycle is higher 
than the pressure needed to keep stream 10 in a liquid phase, stream 10 
may optionally be passed through one or more expansion means, such as a 
hydraulic turbine 40, to produce a PLNG product at a lower pressure but 
still having a temperature above about -112.degree. C. (-170.degree. F.) 
and a pressure sufficient for the liquid product to be at or below its 
bubble point. The PLNG is then sent (stream 20) to a suitable 
transportation or storage means 41 such as a suitable pipeline or carrier 
such as a PLNG ship, tank truck, or rail car. 
FIG. 2 illustrates another embodiment of the invention and in this and the 
embodiments illustrated in FIGS. 1 and 3, the parts having like numerals 
have the same process functions. Those skilled in the art will recognize, 
however, that the process equipment from one embodiment to another may 
vary in size and capacity to handle different fluid flow rates, 
temperatures, and compositions. Referring to FIG. 2, a natural gas feed 
stream enters the system through line 10 and is passed through heat 
exchangers 30 and 31 which are refrigerated by a first refrigeration cycle 
32. Refrigeration cycle 32 cools the feed stream 10 and cools refrigerant 
in a second refrigeration cycle 33 which is further downstream in the 
liquefaction process. 
After exiting the last heat exchanger 31, the feed gas stream 10 enters a 
conventional phase separator 34. A liquid stream 11 exits the bottom of 
the separator and is passed to a conventional demethanizer 35. The 
demethanizer produces an overhead vapor stream 12 which is rich in methane 
and a bottom liquid stream 13 which is predominantly natural gas liquids 
(NGL), primarily ethane, propane, butane, pentane, and heavier 
hydrocarbons. The demethanizer bottoms stream 13 is passed to a 
conventional fractionation plant 36, the general operation of which is 
known to those skilled in the art. The fractionation plant 36 may comprise 
one or more fractionation columns (not shown in FIG. 2) which separate 
liquid bottom stream 13 into predetermined amounts of ethane, propane, 
butane, pentane, and hexane. These liquids are withdrawn from the 
fractionation plant 36 as condensate products, which are collectively 
depicted in the FIG. 2 as stream 14. Overhead streams from the 
fractionation plant 36 are rich in ethane and other light hydrocarbons. 
These overhead streams are collectively shown in FIG. 2 as stream 15. The 
fractionation plant preferably comprises multiple fractionation columns 
(not shown) such as a deethanizer column that produces ethane, a 
depropanizer column that produces propane, and a debutanizer column that 
produces butane, which can be used as make-up refrigerants for the cascade 
refrigeration system (first and second refrigeration cycles 32 and 33) or 
any other suitable refrigeration system. The refrigerant make-up streams 
are collectively illustrated in FIG. 2 by line 16. Although not shown in 
FIG. 2, if feed stream 10 contains high concentrations of CO.sub.2, one or 
more of the refrigerant make up streams may need to be treated to remove 
CO.sub.2 to avoid potential plugging problems in the refrigeration 
equipment. If the CO.sub.2 concentration in the feed stream exceeds about 
3 mole percent, the fractionation plant 36 will preferably include a 
CO.sub.2 removal process. 
The methane-rich stream 17 from the separator 34, the methane-rich stream 
12 from the demethanizer 35, and stream 15 from the fractionation plant 36 
are combined and passed as stream 18 to a series of heat exchangers 37, 
38, and 39 to liquefy the natural gas. Refrigeration to heat exchangers 
37, 38, and 39 are provided by the second refrigeration cycle 33 described 
above. Although the refrigerants in the first and second refrigeration 
cycles 32 and 33 circulate in a closed-loop system, if refrigerants are 
lost from the system through leaks, make up refrigerants can be obtained 
from the fractionation plant 36 (line 16). In the liquefaction process 
illustrated in FIG. 2, only two cycles of a cascade system are needed to 
refrigerate the natural gas stream 10 in accordance with the practice of 
this invention. 
Liquefied natural gas stream 19 exiting the last heat exchanger 39 is 
passed through one or more expansion means, such as hydraulic turbine 40, 
to produce PLNG product at a temperature above about -112.degree. C. 
(-170.degree. F.) and a pressure sufficient for the liquid product to be 
at or below its bubble point. The PLNG is then sent by line 20 to a 
suitable storage means 41. 
In the storage, transportation, and handling of liquefied natural gas, 
there can be a considerable amount of "boil-off," the vapor resulting from 
evaporation of a liquefied natural gas. This invention is particularly 
well suited for liquefying boil-off vapor produced by PLNG. The process of 
this invention can optionally re-liquefy such boil-off vapor. Referring to 
FIG. 2, boil-off vapor may be introduced to the process of the invention 
through line 21. Optionally, a portion of stream 21 may be withdrawn as 
stream 22 and directed through a heat exchanger 42 to cool vapor stream 18 
and to warm the withdrawn boil-off gas for later use as fuel for the 
liquefaction plant. The remaining portion of stream 21 is passed through a 
conventional compressor 43 to compress the boil-off vapor to approximately 
the pressure of vapor stream 18 and is then combined with stream 18. 
FIG. 3 illustrates another embodiment of the present invention. The process 
illustrated in FIG. 3 is similar to the process described above with 
respect to FIG. 2 except that as shown in FIG. 3 stream 18 is passed 
through a compressor 44 and the compressed vapor stream 18 is then passed 
through heat exchangers 45 and 46 which are cooled by refrigerant of the 
first refrigeration cycle 32. 
As illustrated in FIG. 3, boil-off gas may optionally be introduced to 
stream 18 after stream 18 has been cooled by the first refrigeration cycle 
32 and before being cooled by the second refrigeration cycle 33. At least 
a portion of boil-off vapor stream 21 is compressed by a conventional 
compressor 43 and the compressed gas (stream 23) is cooled by a heat 
exchanger 42 which is cooled by stream 22 which has been drawn off from 
stream 21. Stream 22 after being heated by heat exchanger 42 may be used 
as fuel in the liquefaction plant. 
Although FIGS. 2 and 3 show the boil-off vapor being introduced to the 
liquefaction process at a point after fractionation stages and before the 
cooling stages of the second refrigeration cycle, in the practice of this 
invention the boil-off vapor can be introduced to the gas stream to be 
liquefied at any point in the process from before exchanger 30 to after 
exchanger 39 and before expander 40. 
This invention is not limited to any type of heat exchanger, but because of 
economics, plate-fin exchangers and cold box heat exchangers are 
preferred. Preferably all streams containing both liquid and vapor phases 
that are sent to heat exchangers have both the liquid and vapor phases 
equally distributed across the cross section area of the passages they 
enter. To accomplish this, it is preferred to provide distribution 
apparati for individual vapor and liquid streams. Separators can be added 
to the multi-phase flow streams as required to divide the streams into 
liquid and vapor streams. Such separators could be added to the processes 
illustrated in FIGS. 2 and 3 before heat exchangers 38 and 39. 
EXAMPLE 
A simulated mass and energy balance was carried out to illustrate the 
embodiments illustrated in the Figures, and the results are set forth in 
the Tables below. 
The data were obtained using a commercially available process simulation 
program called HYSYS.TM., however, other commercially available process 
simulation programs can be used to develop the data, including for example 
HYSIM.TM., PROII.TM., and ASPEN PLUS.TM., which are all familiar to those 
of ordinary skill in the art. The data presented in Table 1 are offered to 
provide a better understanding of the embodiment shown in FIG. 2, but the 
invention is not to be construed as unnecessarily limited thereto. The 
temperatures and flow rates are not to be considered as limitations upon 
the invention which can have many variations in temperatures and flow 
rates in view of the teachings herein. In this embodiment, the first 
refrigeration cycle 32 is a propane system, and the second refrigeration 
cycle 33 is an ethylene system. 
The data in Table 2 are offered to provide a better understanding of the 
embodiment shown in FIG. 3. In this embodiment, the first refrigeration 
cycle 32 is a propane system, and the second refrigeration cycle 33 is an 
ethane system. 
Using the basic process flow scheme shown in FIG. 1 and using the same feed 
stream composition and temperature, the required total installed power to 
produce conventional LNG (at near atmospheric pressure and a temperature 
of -160.degree. C. (-256.degree. F.) was more than twice the total 
installed power requirement to produce PLNG using the embodiment 
illustrated in FIG. 1: 177,927 kW (238,600 hp) to produce LNG versus 
75,839 kW (101,700 hp) to produce PLNG. This comparison was performed 
using the HYSYS.TM. process simulator. 
A person skilled in the art, particularly one having the benefit of the 
teachings of this patent, will recognize many modifications and variations 
to the specific processes disclosed above. For example, a variety of 
temperatures and pressures may be used in accordance with the invention, 
depending on the overall design of the system and the composition of the 
feed gas. Also, the feed gas cooling train may be supplemented or 
reconfigured depending on the overall design requirements to achieve 
optimum and efficient heat exchange requirements. As discussed above, the 
specifically disclosed embodiments and examples should not be used to 
limit or restrict the scope of the invention, which is to be determined by 
the claims below and their equivalents. 
TABLE 1 
__________________________________________________________________________ 
Phase Flow Rate 
Vapor/ 
Pressure 
Temperature 
kgmole/ 
lbmole/ 
Composition, mole % 
Stream 
Liquid 
kPa 
psia 
.degree. C. 
.degree. F. 
hr hr C.sub.1 
C.sub.2 
C.sub.3+ 
CO.sub.2 
N.sub.2 
__________________________________________________________________________ 
10 V/L 5,516 
800 
4.4 
40 
36,707 
80,929 
92.6 
3.9 
2.48 
0.98 
0.04 
11 L 8,378 
780 
-34.4 
-30 
1,285 
2,833 
38.13 
9.61 
50.97 
1.29 
0 
12 V 5,364 
778 
-34.4 
-30 
473 
1,043 
94.6 
3.69 
0.73 
0.97 
0.01 
13 L 5,378 
780 
187.8 
370 
817 
1,801 
5.43 
13.04 
80.05 
1.48 
0 
14 L 138 
20 
26.7 
80 
553 
1,219 
0 0 100 
0 0 
15 V/L 5,295 
768 
71.7 
161 
224 
494 
19.54 
46.61 
33.85 
0 0 
16 L 3,378 
490 
13.3 
56 
25 
55 
0 2.73 
97.26 
0.01 
0 
17 V 5,378 
780 
-34.4 
-30 
35,422 
78,096 
94.58 
3.69 
0.72 
0.97 
0.04 
18 V 5,295 
768 
-29.4 
-21 
36,120 
79,634 
94.11 
3.96 
0.93 
0.96 
0.04 
19 L 5,019 
728 
-92.8 
-135 
37,469 
82,609 
94.29 
3.84 
0.89 
0.94 
0.04 
20 L 2,861 
415 
-95.6 
-140 
37,469 
82,609 
94.29 
3.84 
0.89 
0.94 
0.04 
21 V 2,827 
410 
-90.0 
-130 
2,724 
6,007 
99.11 
0.46 
0.01 
0.28 
0.14 
22 V 2,827 
410 
-90.0 
-130 
1,375 
3,031 
99.11 
0.46 
0.01 
0.28 
0.14 
__________________________________________________________________________ 
Power 
Power 
Power 
hp kW 
__________________________________________________________________________ 
Compressors 
32, Stage 1 
18,000 
13,423 
32, Stage 2 
35,400 
26,398 
33, Stage 1 
3,300 
2,461 
33, Stage 2 
14,300 
10,664 
33, Stage 3 
29,000 
21,626 
43 450 
336 
36 60 
45 
Expander 
40 -1,200 
-895 
Pump 
36 30 
22 
Net Power Installed 
99,300 
74,049 
Total Installed 
101,700 
75,839 
__________________________________________________________________________ 
TABLE 2 
__________________________________________________________________________ 
Phase Flow Rate 
Vapor/ 
Pressure 
Temperature 
kgmole/ 
lbmole/ 
Composition, mole % 
Stream 
Liquid 
kPa psia 
.degree. C. 
.degree. F. 
hr hr C.sub.1 
C.sub.2 
C.sub.3+ 
CO.sub.2 
N.sub.2 
__________________________________________________________________________ 
10 V/L 5,516 
800 4.4 
40.0 
36,707 
80,929 
92.6 
3.9 
2.48 
0.98 
0.04 
11 L 5,378 
780 -34.4 
-30.0 
1,285 
2,833 
38.13 
9.61 
50.97 
1.29 
0 
12 V 5,364 
778 -34.4 
-30.0 
498 
1,098 
94.61 
3.69 
0.72 
0.97 
0.01 
13 L 5,378 
780 220.0 
428.0 
787 
1,735 
2.35 
13.36 
82.8 
1.49 
0 
14 L 138 
20 26.7 
80.0 
553 
1,219 
0 0 100 
0 0 
15 V/L 5,295 
768 73.9 
165.0 
194 
428 
8.57 
47.09 
38.91 
5.43 
0 
16 L 3,378 
490 13.3 
56.0 
40 
88 
4.52 
32.87 
62.6 
0.01 
0 
17 V 5,378 
780 -34.4 
-30.0 
35,422 
78,096 
94.58 
3.69 
0.72 
0.97 
0.04 
18 V 5,295 
768 -33.3 
-28.0 
36,115 
79,623 
94.11 
3.93 
0.93. 
0.99 
0.04 
19 L 9,997 
1,450 
-87.8 
-126.0 
37,554 
82,796 
94.31 
3.79 
0.89 
0.97 
0.04 
20 L 2,861 
415 -95.6 
-140.0 
37,554 
82,796 
94.31 
3.79 
0.89 
0.97 
0.04 
21 V 2,827 
410 -90.0 
-130.0 
2,724 
6,007 
99.11 
0.46 
0.01 
0.28 
0.14 
22 V 2,827 
410 -90.0 
-130.0 
1,285 
2,833 
99.11 
0.46 
0.01 
0.28 
0.14 
23 V 10,273 
1,490 
-3.3 
26.0 
1,439 
3,173 
99.11 
0.46 
0.01 
0.28 
0.14 
__________________________________________________________________________ 
Power 
Power 
Power 
hp kW 
__________________________________________________________________________ 
Compressors 
32, Stage 1 
15,800 
11,782 
32, Stage 2 
35,100 
26,174 
33. Stage 1 
1,400 
1,044 
33, Stage 2 
7,600 
5,667 
33, Stage 3 
14,800 
11,037 
43 1,100 
820 
44 18,200 
13,572 
36 30 
22 
Expander 
40 -3,900 
-2,908 
Pump 0 
36 30 
22 
Net Power Installed 
90,200 
67,263 
Total Installed 
98,000 
73,080 
__________________________________________________________________________