U.S. patent application number 12/435869 was filed with the patent office on 2010-11-11 for pre-cooled liquefaction process.
This patent application is currently assigned to Air Products and Chemicals, Inc.. Invention is credited to Mark Julian Roberts, Vishal Anandswarup Varma.
Application Number | 20100281915 12/435869 |
Document ID | / |
Family ID | 42362617 |
Filed Date | 2010-11-11 |
United States Patent
Application |
20100281915 |
Kind Code |
A1 |
Roberts; Mark Julian ; et
al. |
November 11, 2010 |
Pre-Cooled Liquefaction Process
Abstract
A system and method for liquefying a natural gas stream, the
method including the steps of providing a dehydrated natural gas
stream for liquefaction, pre-cooling the dehydrated natural gas
stream in a pre-cooling apparatus, where the pre-cooling is
performed by using a pre-coolant that consists essentially of a
hydroflorocarbon (HFC) refrigerant, further cooling the pre-cooled
dehydrated natural gas stream in a main heat exchanger through
indirect heat exchange against a vaporized hydrocarbon mixed
refrigerant coolant to produce a liquefied natural gas product
stream, where the mixed refrigerant coolant comprises ethane,
methane, nitrogen, and less than or equal to 3 mol % of
propane.
Inventors: |
Roberts; Mark Julian;
(Kempton, PA) ; Varma; Vishal Anandswarup;
(Macungie, PA) |
Correspondence
Address: |
AIR PRODUCTS AND CHEMICALS, INC.;PATENT DEPARTMENT
7201 HAMILTON BOULEVARD
ALLENTOWN
PA
181951501
US
|
Assignee: |
Air Products and Chemicals,
Inc.
Allentown
PA
|
Family ID: |
42362617 |
Appl. No.: |
12/435869 |
Filed: |
May 5, 2009 |
Current U.S.
Class: |
62/612 ; 165/166;
62/621 |
Current CPC
Class: |
F25J 1/0256 20130101;
C09K 5/041 20130101; F25J 1/0055 20130101; F25J 1/0087 20130101;
F25J 1/0092 20130101; F25J 1/0292 20130101; F25J 2280/10 20130101;
F25J 2220/64 20130101; F25J 1/025 20130101; F25J 1/0283 20130101;
F25J 1/0022 20130101; F25J 1/0241 20130101; F25J 1/0214 20130101;
F25J 1/0216 20130101; F25J 2290/02 20130101; F25J 1/0278 20130101;
F25B 9/006 20130101; F25J 1/004 20130101; F25J 1/0215 20130101;
F25J 1/0298 20130101; F25J 1/0052 20130101; F25J 1/021 20130101;
F25J 1/0097 20130101; F25J 1/0284 20130101; F25J 1/0238
20130101 |
Class at
Publication: |
62/612 ; 62/621;
165/166 |
International
Class: |
F25J 1/00 20060101
F25J001/00; F25J 3/06 20060101 F25J003/06; F28F 3/00 20060101
F28F003/00 |
Claims
1. A method for liquefying a natural gas stream, the method
comprising the steps of: providing a dehydrated natural gas stream
for liquefaction; pre-cooling the dehydrated natural gas stream in
a pre-cooling apparatus, wherein the pre-cooling is performed by
using a pre-coolant that consists essentially of a hydroflorocarbon
(HFC) refrigerant; further cooling the pre-cooled dehydrated
natural gas stream in a main heat exchanger through indirect heat
exchange against a vaporized hydrocarbon mixed refrigerant coolant
to produce a liquefied natural gas product stream, wherein the
mixed refrigerant coolant comprises ethane, methane, nitrogen, and
less than or equal to 3 mol % of propane.
2. The method of claim 1, wherein the method for liquefying the
natural gas stream occurs on a Floating Production Storage and
Offloading platform (FPSO).
3. The method of claim 1, wherein the mixed refrigerant coolant
comprises less than 2 mol % of propane.
4. The method of claim 1, wherein the mixed refrigerant coolant
comprises less than 1 mol % of propane.
5. The method of claim 1, wherein the hydroflorocarbon refrigerant
is R410A.
6. The method of claim 1, wherein the hydroflorocarbon refrigerant
has a temperature glide of less than or equal to 7.degree. C.
7. The method of claim 1, further comprising feeding the pre-cooled
dehydrated natural gas stream to a scrub column where the
pre-cooled dehydrated natural gas stream is rectified and stripped
of a portion of the heavy hydrocarbons present in the pre-cooled
dehydrated natural gas stream, feeding the rectified and stripped
natural gas stream from the scrub column back into the pre-cooling
apparatus where the rectified and stripped natural gas stream is
further cooled and partially condensed, separating in a separator
the cooled and partially condensed natural gas stream, where the
liquid portion of the natural gas stream is pumped and then sent to
the scrub column as a reflux stream, and where the vapor portion of
the natural gas stream is sent to the main heat exchanger to
produce the liquefied natural gas product stream.
8. The method of claim 1, further comprising pre-cooling the
hydrocarbon mixed refrigerant coolant in the pre-cooling apparatus
prior to vaporization in the main heat exchanger.
9. The method of claim 1, further comprising developing pressure in
an air compressor driven by a gas generator prior to starting a
power turbine, wherein the power turbine drives a multi-stage HFC
compressor for pre-cooling of at least the dehydrated natural gas
stream for liquefaction; and using a high temperature, high
pressure gas discharged from the gas generator to provide the
required high starting torque to the power turbine, wherein the
starting of the power turbine occurs without venting of the HFC
refrigerant.
10. A system for liquefying a natural gas stream, comprising: a
multistage hydroflorocarbon (HFC) compressor; a pre-cooling
apparatus fluidly connected to the multistage HFC compressor, the
pre-cooling apparatus comprising at least one evaporator for
pre-cooling a dehydrated natural gas stream using a HFC
refrigerant; a main heat exchanger fluidly connected to the
pre-cooling apparatus for further cooling the pre-cooled dehydrated
natural gas stream to produce a liquid natural gas product stream,
wherein a vaporized mixed refrigerant coolant comprising ethane,
methane, nitrogen, and less than or equal to 3 mol % of propane
provides refrigeration in the main heat exchanger.
11. The system of claim 10, wherein the system operates on a
Floating Production Storage and Offloading platform (FPSO).
12. The system of claim 10, wherein the vaporized mixed refrigerant
coolant comprising less than or equal to 2 mol % of propane
provides refrigeration in the main heat exchanger.
13. The system of claim 10, further comprising a multi-shaft gas
turbine driver for supplying the power to the multistage HFC
compressor, the multi-shaft gas turbine driver comprising: a power
turbine connected to the multistage HFC compressor on a first
shaft; a gas generator fluidly connected to the power turbine on a
second shaft; an air compressor connected to the gas generator on
the second shaft; and a combustor fluidly connected to the air
compressor and the gas generator for supplying a high temperature,
high pressure gas to the gas generator.
14. The system of claim 10, further comprising an electrical motor
assembly for supplying the power to the multistage compressor, the
electrical motor assembly comprising: an electrical motor connected
to the multistage HFC compressor; a power supply connected to the
electrical motor.
15. The system of claim 14, wherein the power supply of the
electrical motor is a power grid of a Floating Production Storage
and Offloading platform (FPSO) or an electrical generator driven by
a multi-shaft gas turbine.
16. The system of claim 10, wherein the evaporator is a
shell-and-tube type evaporator.
17. The system of claim 10, wherein the main heat exchanger is a
wound coil exchanger.
18. The system of claim 10, further comprising a conduit between
the main heat exchanger and the pre-cooling apparatus to feed the
mixed refrigerant coolant to the pre-cooling apparatus prior to
vaporization of the mixed refrigerant coolant in the main heat
exchanger.
19. A method for liquefying a natural gas stream, the method
comprising the steps of: providing a dehydrated natural gas stream
for liquefaction; pre-cooling the dehydrated natural gas stream in
a pre-cooling apparatus, wherein the pre-cooling is performed by
using a pre-coolant that consists essentially of a hydroflorocarbon
(HFC) refrigerant; further cooling the pre-cooled dehydrated
natural gas stream in a main heat exchanger through indirect heat
exchange against a vaporized hydrocarbon mixed refrigerant coolant
to produce a liquefied natural gas product stream, wherein the
mixed refrigerant coolant is a mixture comprising a methane stream
derived from a natural gas stream, an ethane enriched stream,
wherein the ethane enriched stream is predominantly ethane, and a
nitrogen stream, wherein the nitrogen stream is predominantly
nitrogen.
20. The method of claim 19, wherein the method for liquefying the
natural gas stream occurs on a Floating Production Storage and
Offloading platform (FPSO).
Description
BACKGROUND
[0001] Previously pre-cooled natural gas liquefaction processes
were disclosed for use on Floating Production Storage and
Offloading platforms (FPSO's) that used CO.sub.2 to pre-cool a
natural gas feed stream while a mixed refrigerant hydrocarbon
mixture (HMR) was used to further cool the pre-cooled stream to
provide a liquefied natural gas (LNG) product. In these processes,
the natural gas feed stream is pre-cooled against boiling CO.sub.2
at one or more pressure levels. The CO.sub.2 vaporizes while
pre-cooling the natural gas feed stream to a temperature of
approximately -35.degree. C. The CO.sub.2 vapors are then
compressed, cooled, and condensed to form the liquid CO.sub.2
refrigerant to be re-circulated back into the process.
[0002] Using CO.sub.2 as a precoolant for liquefaction of LNG on a
FPSO, however, has several disadvantages. First, CO.sub.2 has a
freezing point of -56.6.degree. C. at which temperature dry ice
formation begins to occur. To prevent operational issues associated
with the formation of dry ice in a liquefaction plant, it has been
suggested that the high pressure CO.sub.2 preferably not be cooled
below -40.degree. C. Assuming that the high pressure CO.sub.2 is
cooled to -40.degree. C. and a temperature approach of at least
3.degree. C., a natural gas feed cannot be cooled below -37.degree.
C. when using CO.sub.2 as a precoolant. Often, however, the natural
gas feed must be pre-cooled to temperatures below -37.degree. C. in
order to condense and remove heavy hydrocarbons and aromatics with
the goal of reducing the heating value of the LNG product and/or
preventing freeze-out of impurities such as benzene in the LNG
product. Use of CO.sub.2, therefore, may preclude the removal of
required amounts of such impurities in the pre-cooling section.
[0003] Second, the critical temperature (i.e., the temperature
above which a fluid cannot be condensed irrespective of pressure)
of CO.sub.2 is approximately 31.1.degree. C., which means that
CO.sub.2 cannot be condensed above 31.1.degree. C. A pre-cooling
cycle, however, requires a condensed refrigerant which supplies
refrigeration by boiling against a load stream. Traditionally
FPSO's use sea water as a coolant supplied to compressor
inter-coolers and after-coolers. Thus, if the sea water is warm,
for example, 27.degree. C., which is the typical surface water
temperature in the tropical regions, and assuming a typical
10.degree. C. approach in the condenser, the CO.sub.2 cannot be
condensed using the mild sea water, thereby degrading the
efficiency and limiting the applicability of CO.sub.2 based cycles
for latitudes with sea water below approximately 20.degree. C.
[0004] Additionally, the CO.sub.2 needs to be typically compressed
to a pressure above approximately 52 bara for it to be efficiently
condensed, which requires the use of special high pressure barrel
type casing for the compressor. The use of special high pressure
barrel type casings thereby raises the capital cost. Moreover, the
overall pre-cooling loop works at much higher pressures than
propane or hydrofluorocarbon based pre-cooling cycles. Thus, the
whole system must be designed with high pressure piping, pressure
relief devices, etc. that increase capital cost and increase gas
pressure safety concerns, especially on an offshore platform where
distances between equipment and personnel is much less than
compared with land-based plants.
[0005] Pre-cooled natural gas liquefaction systems and processes
for FPSO's that use propane (C.sub.3H.sub.8) to pre-cool a natural
gas feed stream while a mixed refrigerant hydrocarbon mixture,
which itself contains propane, further cools the pre-cooled stream
to provide a liquefied natural gas product have also been
disclosed. This type of process is known as the propane pre-cooled
mixed refrigerant, or C3MR process, and is used to manufacture most
of the LNG produced worldwide. In the pre-cooling loop, liquid
propane is expanded to different pressures using Joule-Thomson
(J-T) valves. The resulting boiling propane vaporizes against the
natural gas feed stream to pre-cool the natural gas feed stream.
The resulting propane vapors are fed to a propane compressor that
compresses the vapor streams to a high pressure. The high pressure
propane discharging from the compressor is cooled against sea water
and re-circulated back into the pre-cooling process. Due to the
high critical temperature of propane (96.6.degree. C.) versus
CO.sub.2 (31.1.degree. C.), propane provides high refrigeration
duties for every megawatt of supplied compression power in
comparison to CO.sub.2. In the classic land-based plants, propane
may also used in the mixed hydrocarbon refrigerant in order to
provide efficient refrigeration in the liquefaction part of the
process.
[0006] While propane has been widely used in land-based plants as a
pre-cooling refrigerant, use of propane also has disadvantages for
use on LNG FPSOs. Propane may leak from gas compressor seals,
propane evaporators, and other points in the system that can lead
to hazardous explosive conditions at and near the surface of the
FPSO. Propane presents these hazardous explosive conditions due to
its several unique properties. For example, propane has a normal
boiling point of -42.1.degree. C., which means that when propane
leaks from any equipment, it remains a vapor. Propane's vapor
density is 1.91 Kg/m3 versus a density of air of 1.20 Kg/m3 at
15.degree. C. Propane's vapor density causes the leaked vapor to
settle near the surface of the FPSO platform. The vapor has a low
flammability limit of less than 9.5 volume % and the U.S.
Department of Transportation lists propane as a flammable gas,
while the European Commission labels propane as F+ or extremely
flammable. Because the propane vapors settle to the floor or
surface of the FPSO platform and because propane is extremely
flammable, the presence of a propane "cloud" on the FPSO platform
can be extremely hazardous.
[0007] Moreover, an FPSO is space-constrained and the average
distance between equipment and personnel or living quarters is
significant smaller than for a land-based plant. If a propane leak
were to occur on an FPSO, the flammable propane vapors would likely
propagate to all major areas of the FPSO in a short span of time,
exposing the plant to explosion and/or fire risk.
[0008] This hazard potential became a reality on the Piper Alpha,
an oil & gas platform processing propane condensate in the
North Sea. On Jul. 6, 1988, a back-up propane condensate pump which
was allegedly under repair was accidentally started which led to
the discharge of propane into the air just above the surface of the
platform. The resultant leaked propane was accidentally ignited and
resulted in two very large explosions that immediately engulfed the
control room and most of the ship. The accident resulted in 167
fatalities and the platform melted.
[0009] Use of propane as a refrigerant can, therefore, be extremely
hazardous, especially since it will require propane inventories not
just in the pre-cooling section but also in storage and in a
de-propanizer column to extract the propane. In contrast, other
hydrocarbons do not pose as great a threat as propane. Methane
(CH.sub.4) is, for example, lighter than air with a typical vapor
density of 0.68 Kg/m3 at 15.degree. C. If methane were to leak on a
FPSO platform, the methane would quickly dissipate in the air.
Similarly, ethane (C.sub.2H.sub.6) has a density of 1.28 Kg/m3 at
15.degree. C., which is still close to the density of air, thus,
reducing the hazard of flammable vapors gathering close to the FPSO
surface.
[0010] Butanes (C.sub.4H.sub.10) typically have similar issues
associated with propane. N-Butane, for example, has a normal
boiling point of -0.5.degree. C. and a density of 2.52 Kg/m3 at
15.degree. C. that could lead to the formation of explosive vapors
on the FPSO surface. Butane, however, has a lower tendency to
escape the equipment of the liquefaction plant due to its high
density, which renders it inherently safer than propane. Heavier
hydrocarbons (C5+) are primarily liquids at ambient atmospheric
temperatures, and hence, any leaks will result in fewer vapors
close to the FPSO surface.
[0011] Thus, due to its unique position amongst the group of
aliphatic hydrocarbons, propane is considered the most hazardous of
all aliphatic hydrocarbons, and it is desirable to avoid for use on
LNG FPSOs.
[0012] There is, therefore, a need in the art for a process that
provides the energy and capital cost efficiency of the C3MR process
without exposing, or at least limiting the use of flammable
refrigerants on the FPSO. A pre-cooling refrigerant is needed that
will have a critical temperature higher than 50.degree. C., will
have zero flammability potential, zero toxicity potential, low
environmental impact, and a low normal boiling point that allows
low pre-cooling temperatures.
BRIEF SUMMARY
[0013] Embodiments of the present invention satisfy this need in
the art by providing an improved liquefaction system and method
that utilizes a hydrofluorocarbon (HFC) to pre-cool a natural gas
feed stream prior to liquefying the pre-cooled feed using a mixed
refrigerant (MR).
[0014] In one embodiment, a method for liquefying a natural gas
stream is disclosed, where the method comprises the steps of:
providing a dehydrated natural gas stream for liquefaction;
pre-cooling the dehydrated natural gas stream in a pre-cooling
apparatus, wherein the pre-cooling is performed by using a
pre-coolant that consists essentially of a hydroflorocarbon (HFC)
refrigerant; further cooling the pre-cooled dehydrated natural gas
stream in a main heat exchanger through indirect heat exchange
against a vaporized hydrocarbon mixed refrigerant coolant to
produce a liquefied natural gas product stream, wherein the mixed
refrigerant coolant comprises ethane, methane, nitrogen, and less
than or equal to 3 mol % of propane.
[0015] In another embodiment, a system for liquefying a natural gas
stream is disclosed, where the system comprises: a multistage
hydroflorocarbon (HFC) compressor; a pre-cooling apparatus fluidly
connected to the multistage HFC compressor, the pre-cooling
apparatus comprising at least one evaporator for pre-cooling a
dehydrated natural gas stream using a HFC refrigerant; a main heat
exchanger fluidly connected to the pre-cooling apparatus for
further cooling the pre-cooled dehydrated natural gas stream to
produce a liquid natural gas product stream, wherein a vaporized
mixed refrigerant coolant comprising ethane, methane, nitrogen, and
less than or equal to 3 mol % of propane provides refrigeration in
the main heat exchanger.
[0016] In yet another embodiment, a method for liquefying a natural
gas stream is disclosed where the method comprises the steps of:
providing a dehydrated natural gas stream for liquefaction;
pre-cooling the dehydrated natural gas stream in a pre-cooling
apparatus, wherein the pre-cooling is performed by using a
pre-coolant that consists essentially of a hydroflorocarbon (HFC)
refrigerant; further cooling the pre-cooled dehydrated natural gas
stream in a main heat exchanger through indirect heat exchange
against a vaporized hydrocarbon mixed refrigerant coolant to
produce a liquefied natural gas product stream, wherein the mixed
refrigerant coolant is a mixture comprising a methane stream
derived from a natural gas stream, an ethane enriched stream,
wherein the ethane enriched stream is predominantly ethane, and a
nitrogen stream, wherein the nitrogen stream is predominantly
nitrogen.
BRIEF DESCRIPTION OF THE DRAWINGS
[0017] The foregoing brief summary, as well as the following
detailed description of exemplary embodiments, is better understood
when read in conjunction with the appended drawings. For the
purpose of illustrating embodiments of the invention, there is
shown in the drawings exemplary constructions of the invention;
however, the invention is not limited to the specific methods and
instrumentalities disclosed. In the drawings:
[0018] FIG. 1 is a flow chart illustrating an exemplary system and
method involving aspects of the present invention;
[0019] FIG. 2 is a flow chart illustrating an exemplary system and
method involving aspects of the present invention;
[0020] FIG. 3 is a flow chart illustrating an exemplary system and
method involving aspects of the present invention;
[0021] FIG. 4A is a flow chart illustrating an exemplary driver
system and method involving aspects of the present invention;
and
[0022] FIG. 4B is a flow chart illustrating an exemplary driver
system and method involving aspects of the present invention.
DETAILED DESCRIPTION
[0023] One embodiment of the invention concerns the development of
cryogenic refrigeration for use in LNG liquefaction applications,
and especially for use on offshore platforms based upon the use of
both HFC's and hydrocarbon refrigerant mixtures. Safety is of
paramount importance in the design of any offshore or floating
hydrocarbon processing plant due to the compact layout, proximity
of the living quarters to the process equipment, and the limited
egress. Embodiments of the disclosed system and process achieve
high power efficiency per ton of product LNG, thus, eliminating the
need for using flammable, but highly efficient refrigerants like
propane.
[0024] In one embodiment, the highly flammable propane pre-cooling
refrigerant used in the traditional C3MR process may be replaced
with an HFC to pre-cool the natural gas feed before being further
cooled using a hydrocarbon mixed refrigerant. The HFC used may be
classified (class 1) by the American Society of Heating,
Refrigerating and Air Conditioning Engineers (ASHRAE) as being
non-toxic under 400 ppm and does not propagate a flame at
21.degree. C. and 1.01 bara, thus, rendering it safe for use in
comparison to propane. ASHRAE classifies a refrigerant to be type 2
if it is flammable at concentrations greater than 0.1 Kg/m3 at
21.degree. C. and 1.01 bara with a heat of combustion less than
19,000 KJ/Kg. ASHRAE classifies a refrigerant to be type 3 if it is
flammable at concentrations below 0.1 Kg/m3 at 21.degree. C. and
1.01 bara with a heat of combustion greater than 19,000 KJ/Kg.
[0025] The HFC will allow pre-cooling to temperatures lower than
-37.degree. C. due to a lower normal boiling point for many of the
HFC's. For example, R410A which is a binary mixture of R32
(difluoromethane) and R125 (pentafluoroethane) has a critical
temperature of 70.1.degree. C., classified as an Al refrigerant by
ASHRAE (low toxicity and no flammability potential), has an ozone
depletion potential (ODP) of zero, and a normal boiling point of
-51.6.degree. C. The ODP is the tendency of a molecule to react and
destroy atmospheric ozone relative to dichlorodifluromethane
(CCl.sub.2F.sub.2) which is assigned an ODP of 1.0 as per the
provisions of the Montreal protocol (1987). Refrigerants with a
lower ODP are, thus, more desirable.
[0026] Another benefit of using HFC's as a pre-cooling refrigerant
is that it allows removal of propane from the mixed refrigerant
maintaining close to the same energy efficiency of producing LNG.
This allows for an HFC pre-cooled mixed refrigerant process that
eliminates the use of propane, offering about the same efficiencies
as a land based plant, and at the same time, allowing a high amount
of operating flexibility in the liquefaction process due to the low
boiling temperatures of HFC's.
[0027] Use of an HFC refrigerant can lead to efficiencies as high
as or even higher than the traditional C3MR process. Further,
several mainstream commercially available HFC's allow pre-cooling
to temperatures as low as -127.degree. C. without any possibility
of freeze-out. Moreover, the critical temperature of most
commercially known HFC's is much higher than CO.sub.2, and hence
HFC's may be used with high process efficiency under most typical
sea water temperature conditions.
[0028] Another benefit of using HFC's over propane is the ability
of several HFC's to be flashed down to vacuum (sub-atmospheric)
pressure levels to achieve very low pre-cooling temperatures, thus,
having the ability to condense enough hydrocarbons from the feed.
If propane is flashed down to sub-atmospheric levels to achieve low
temperatures (i.e., below -42.degree. C.), air ingress into the
compressor suction line, for example, can potentially lead to
explosive mixtures. Since HFC's like R410A and R134A, for example,
are non-flammable, air ingress will not lead to flammable mixtures,
which is why it is acceptable to flash down to sub-atmospheric
levels.
[0029] Due to the feasibility of reaching colder pre-cooling
temperatures, it is also possible, or nearly possible, to
eliminate, the propane from the main mixed refrigerant loop while
still maintaining an acceptable commercial efficiency. The mixed
refrigerant used for liquefaction may comprise nitrogen, methane,
and ethane, for example. Theoretically, the mixed refrigerant
should contain zero mole % propane, however, practically,
elimination of all propane in the mixed refrigerant may not be
economical or commercially feasible.
[0030] The methane for the mixed refrigerant may be produced
on-site. The ethane used for preparing the mixed refrigerant may be
sourced from commercial vendors, for example, or may also be
prepared on-site. Whether sourced externally or produced on-site,
ethane is typically produced as the overhead product from a
de-ethanizer column. In practice, the purity of ethane from a
de-ethanizer column is limited by the number of stages in the
column and the reflux ratio. Production of high purity ethane
requires a large number of stages in the distillation column and/or
a high reflux ratio leading to increases in capital and operating
costs. Thus, in practice a small amount of propane may be included
in a commercial ethane product stream.
[0031] The methane required for the mixed refrigerant is typically
extracted as a vapor stream (see stream 117 of FIG. 1) emerging
from a scrub column reflux drum (see item 110 from FIG. 1). The
propane content in the methane stream 117 is a function of the
temperature of the pre-cooled feed stream 108 (referring to FIG.
1). As the temperature of the pre-cooled stream 108 decreases, more
propane condenses in the reflux drum 110 that leads to lower
propane content in overhead streams 117 and 114. Hence, if lower
propane content is desired in the methane stream used to prepare
the mixed refrigerant, then a lower pre-cooling temperature is
needed. Methane can also be procured from the dehydrated feed gas.
In either case, the methane makeup stream will also typically
contain some ethane and propane as impurities, depending on the
plant operation and the composition of the natural gas feed.
[0032] Thus, both sources of methane and ethane contain small
quantities of propane. For example, as per FIGS. 2-3 of the Gas
Processors Supplier Association (GPSA) Engineering Data Book (FPS
Version, vol. I, 11.sup.th ed., 1998) commercially available high
purity ethane commonly contains about 4 mole % propane and may
contain up to 10.7 mol % propane. Table 1 shows the variation of
propane content in a typical mixed refrigerant (where the mixed
refrigerant comprises 16.6 mole % N.sub.2, 40.99 mole % methane,
and 42.4 mole % of ethane). For example, if less than 2.0 mole % of
propane is desired in the mixed refrigerant for safety or other
purposes, then either a methane stream of less than 1.0 mole %
propane would have to be used or the ethane stream would need to
contain less than 4 mole % propane. In this case, ethane from a
high purity de-ethanizer column may be preferred over the
commercial grade ethane.
TABLE-US-00001 TABLE 1 Mole % Propane in Mixed Mole % Refrigerant
Propane in Using Mixed Commercial Refrigerant High Purity Mole %
Using On-site Ethane Propane in Ethane With Containing 4% Methane
1% Propane Propane 0.50 0.63 1.90 1.00 0.83 2.11 1.50 1.04 2.31
2.50 1.45 2.72 2.66 1.51 2.79 3.00 1.65 2.93 3.19 1.73 3.00
[0033] Pre-cooling to lower temperatures reduces the heat removal
load on the mixed refrigerant loop. This results in both capital
and operating cost reductions; mixed refrigerant loop equipment
such as cryogenic heat exchangers and mixed refrigerant compressors
are generally more capital intensive than pre-cooling equipment.
This is because the mixed refrigerant loop heat exchangers are more
elaborately designed to withstand cryogenic conditions and also
because mixed refrigerant compressors operate at several times
higher pressures than pre-cooling compressors leading to thicker
walled piping and equipment. Hence, reducing the load on the mixed
refrigerant loop leads to net capital cost savings. Further, the
mixed refrigerant loop requires higher energy consumption from the
compressors for every KW of heat removed versus the pre-cooling
loop. This is because, the average compressibility factors of
pre-cooling refrigerants like HFC's and propane is much lower than
that of a typical mixed refrigerant. Hence, shifting heat removal
load to the pre-cooling loop provides net reduction in total
compression power leading to operating cost benefits. Propane does
not, however, provide the flexibility to shift duty to the
pre-cooling loop due to its relatively high normal boiling point
(-42.4.degree. C.) in comparison to HFC's like R410A (-51.6.degree.
C.), and the hazard associated with vacuum operation. Elimination,
or near elimination of propane as a refrigerant everywhere in the
plant, along with any associated equipment costs and hazards of
storing and producing propane, will lead to a significantly safer
and more economical liquefaction system.
[0034] HFC refrigerants that are either single component
refrigerants or refrigerant mixtures may be used for pre-cooling.
Single component HFC refrigerants, such as R134a, have been widely
used in the automobile and other refrigeration industries, for
example. Only a few single component HFC refrigerants, however,
satisfy the multiple constraints of being stable, being low
flammable, having low ozone depletion potentials of less than 0.1,
and having a low global warming potential with efficient
thermodynamic characteristics of low condensing pressures and low
normal boiling points. Hence, mixtures are widely considered as
potential candidates for HFC refrigerants.
[0035] A wide range of commercial HFC's are available allowing high
flexibility in setting pre-cooling temperatures. For example, if
the LNG has to be stripped of heavier hydrocarbons (often to
control its heating value), the feed may need to be pre-cooled to
temperatures below -40.degree. C. Commercially available HFC's like
R-410A can be readily used to pre-cool the feed to temperatures of,
for example, -48.degree. C., at very high efficiencies. In
contrast, CO.sub.2 could not be used as a pre-coolant in this
situation because of the possibility of freeze-out.
[0036] HFC refrigerant mixtures considered to have low "temperature
glide" varying between 0-7.degree. C., may be used as the
pre-coolant. The temperature glide is the temperature difference
between the bubble and the dew point temperatures and is due to the
change in composition as a mixture boils gradually moving towards
heavier components as the more volatile components boil out.
Refrigerant mixtures with very small temperature glides
(0-0.1.degree. C.) act like single components and are termed
"azeotropic mixtures" while mixtures with low temperature glides
(i.e., less than 1.degree. C.) have been termed "near-azeotropes."
For example, The HFC mixture R-410A has a temperature glide of less
than 0.1.degree. C.
[0037] Low temperature glide refrigerants have several benefits.
First, low temperature glide refrigerants avoid composition change
so that the whole refrigeration circuit, including the equipment,
may be designed for a uniform composition. In addition, leakage of
a low or zero temperature glide refrigerant will result in minimal
or zero change in the composition of the refrigerant contained in
the pre-cooling system.
[0038] Second, near-azeotropes or azeotropes refrigerants behave
differently from their individual components. For example, R-410A
which is a 50/50 by weight near-azeotropic mixture of R32 and R125,
has a low flammability, even though R32 by itself is flammable.
[0039] Third, an azeotrope refrigerant will maintain its molecular
weight as the mixture is evaporated at various pressure levels.
Thus, vapors entering a compressor at different pressure levels
have the same molecular weight. In contrast, using HFC's with large
temperature glides may lead to vapors with higher molecular weights
to enter the compressor, which could lead to unpredictable behavior
of the compressor unless robust composition control schemes are
deployed.
[0040] One point that must be addressed when pre-cooling with an
HFC refrigerant on a FPSO is the need for full pressure re-starting
of the HFC refrigerant compressor. Many land-based LNG plants use
single shaft "frame" type gas turbines to drive the refrigerant
compressors in conjunction with large electric starter/helper
motors. Even with large starter/helper motors, these plants still
require a large percentage of the refrigerant inventory to be
vented before re-starting the compressor. This is necessary to
reduce the power/torque requirements of the turbine and
starter/helper motor to a level low enough to start the drive
train. The drive train usually consists of one or more compressors
and a driver which are all mechanically connected. Venting is not
an acceptable economic option with an HFC pre-cooled LNG plant,
particularly one located offshore, because the refrigerant is
expensive to purchase and transport, and storage of the HFC would
take up valuable space on an FPSO. To meet the requirements of a
full pressure start without venting, the driver has to be capable
of providing the torque necessary to accelerate the drive train to
full speed under load. This can be done with an electric motor
drive, a multi-shaft frame, or multi-shaft aero derivative gas
turbine as they all have adequate torque capability. The
requirement of not venting the HFC's in an LNG liquefaction system
using HFC pre-cooling cannot likely be met with a single shaft gas
turbine based drive train without an exceptionally large
starter/helper motor and the power generation equipment required to
power it. This configuration would not be an economical solution
since it would require a large investment in equipment that would,
at most, be used occasionally when the plant is re-started.
[0041] FIG. 1 illustrates a liquefaction system and process
including two refrigeration circuits according to one aspect of the
current invention. One or more refrigeration circuits may be added
to the circuits already illustrated in FIG. 1. The first, warm, or
pre-cooling refrigeration circuit utilizes an HFC refrigerant. The
second, cool, or main cooling refrigeration circuit utilizes a
hydrocarbon mixture. The first refrigeration circuit using HFC as a
pre-cooling refrigerant comprises a multi-stage HFC compressor 158
and a HFC pre-cooling apparatus 101. The HFC used in the first
refrigeration circuit may preferably have the following properties:
(1) be an ASHRAE Class A (low toxicity) refrigerant; (2) be an
ASHRAE Class 1 (low flammability) refrigerant; (3) have an ODP not
exceeding 0.1 (very low ozone depleting). The ODP of a molecule
compares its ozone destroying tendency with that of
dichlorodifluromethane (CCl.sub.2F.sub.2) as per the provisions of
the Montreal protocol; and (4) have as low a Global Warming
Potential (GWP). It should be noted, however, that the extremely
low likelihood of large scale emissions under normal operating
conditions, the energy savings, and the consequent reductions in
CO.sub.2 emissions imply that the GWP criterion is only
qualitative.
[0042] The HFC used in the first refrigeration circuit may also
preferably have the following properties: (1) for HFC mixtures, the
refrigerant preferably would have a temperature glide of less than
7.degree. C. which is satisfied by azeotropic or near-azeotropic
HFC mixtures; (2) the HFC component or mixture would have a
critical temperature higher than 50.degree. C. (allowing use of
ambient air as coolant in addition to sea water, thus, imparting a
higher efficiency); and (3) having normal boiling points that are
low enough to pre-cool the natural gas feed to the desired
temperature. For example, R410A has a normal boiling temperature of
-51.6.degree. C. even without vacuum operation, which may allow for
sufficient removal of heavy hydrocarbons from the natural gas feed
using fractionation.
[0043] Commercially available single component HFC's like R134a,
R125, and all others bearing the above properties qualify for use
in the current invented process for LNG liquefaction. Also
commercially available HFC mixtures like R407C, R410A, R417A, R507,
and R422D satisfy the above requirements.
[0044] Table 2 provides a summary of other potential pure and
mixture HFC's that are classified as Al (i.e., no toxicity below
400 ppm and no flame propagation potential), with close to zero
ODP, and that can be used for LNG pre-cooling service.
TABLE-US-00002 TABLE 2 Critical ASHRAE Refrigerant Temp Safety Type
Composition ODP (.degree. C.) Classification R134a
CF.sub.3CH.sub.2F 0 100.9 A1 R125 CF.sub.3CHF.sub.2 0 66.1 A1 R407C
R32/R125/R134a 0 87.3 A1 R410A R32/R125 0 70.1 A1 R507 R143A/R125 0
70.9 A1
[0045] Returning to FIG. 1, a natural gas feed stream (not shown)
is pre-treated for removal of heavy hydrocarbon oils, particulates,
CO.sub.2, and H.sub.2S before being sent to driers (not shown).
Drying may be performed using sea water cooling if the sea water is
substantially below 22.degree. C. or can be performed using the HFC
refrigerant. After cooling the natural gas feed stream to a
temperature between 22-25.degree. C., the natural gas feed stream
is then sent to drier beds where moisture is removed (not shown).
The dehydrated natural gas feed stream 100 is then sent to be
pre-cooled at pressures ranging between 30-85 bara. Pre-cooling of
dehydrated natural gas feed stream 100 is performed in 1-5 cooling
stages in series, for example, represented by the pre-cooling
apparatus 101. FIG. 1 illustrates a 3-stage pre-cooling system.
These serial cooling stages use an HFC refrigerant at sequentially
descending temperatures by lowering J-T valve pressures making the
HFC refrigerant supplied to the cooling stage (n) colder than that
supplied to the cooling stage (n-1), for example. The greater the
number of cooling stages, the greater the efficiency of pre-cooling
due to close approaches of the cooling curve. If there are a total
of (n) HFC pre-cooling stages, then the feed cools in (n-1) stages
to yield the pre-cooled stream 102.
[0046] Pre-cooled stream 102 may then be sent to a hydrocarbon
scrub column 103 which scrubs away heavier (C.sub.3+) components of
the feed using a cold liquid reflux stream 113 in order to adjust
the heating value of the final LNG. A bottoms stream 105 is sent
either to a fractionation train or to storage (not shown). It
should be noted that due to space constraints on FPSOs, the heavy
hydrocarbon stream 105 exiting the scrub column 103 may be
potentially shipped and fractionated at a LNG receiving terminal.
If fractionation is undertaken on the FPSO platform, one aspect of
the current invention also allows for the HFC refrigerant to supply
refrigeration to condensers of the various columns (such as a
de-ethanizer) that may be involved in a fractionation train.
[0047] Stream 104, taken from the scrub column 103, constitutes the
lighter overhead stream. Part of stream 104 (i.e., stream 107) may
be partially condensed using the HFC pre-cooling apparatus 101. The
partially condensed feed stream 108 may then be combined with the
uncondensed portion of stream 104 (i.e., stream 106) to form stream
109 and then sent to a vapor-liquid separator 110 which disengages
the vapor from the liquid. The liquid stream 111 from the
vapor-liquid separator 110 may then be pumped in pump 112 back into
scrub column 103 as stream 113 to act as the column reflux.
[0048] The HFC pre-cooling refrigerant may be used to supply all of
the scrub column reflux condenser 110 duty without the need to use
the main liquefaction refrigerant for such purpose. Using the HFC
pre-cooling to supply all of the scrub column reflux condenser 110
duty will improve the efficiency of the system since typically
cooling duties supplied by the typical hydrocarbon refrigerants
require much higher incremental compression power than the HFC
refrigerant. This is because of the significantly lower
compressibility factors of typical HFC's when compared with lighter
hydrocarbon refrigerants like CH.sub.4 and C.sub.2H.sub.6. Use of
the HFC pre-cooling to supply all of the scrub column reflux
condenser 110 duty also reduces the size of the main liquefaction
exchanger 115 and simplifies control issues and plant layout.
[0049] Vapor stream 114 from the scrub column reflux condenser 110
may be sent to the cryogenic section of the plant that fully
condenses and sub-cools vapor stream 114 to form LNG product stream
116. The cryogenic section comprises the main liquefaction
exchanger 115. In the cryogenic section, either a refrigerant
consisting of mixed hydrocarbons with 0-30 mole % N.sub.2 or pure
N.sub.2 may be used, for example. In one embodiment, the main
liquefaction refrigerant may be a mixture containing 0-30% N.sub.2
and hydrocarbons such as methane (0-50%), ethane (0-75%), and
butanes (0-50%). In another embodiment, the main liquefaction
refrigerant may be a mixture comprising a first stream of methane
derived from a natural gas stream, a second stream, where the
second stream is an ethane enriched stream that is predominantly
ethane, and a third stream, where the third stream is a nitrogen
enriched stream that is predominantly nitrogen. The methane stream
can be derived from natural gas in one of two ways. If natural gas
stream 100 (illustrated in FIG. 1) is lean (i.e., contains more
than 90 mole % methane and less than 3 mole % propane) then a part
of that stream may be used to make up the mixed refrigerant. If
natural gas stream 100 (of FIG. 1) is not lean (i.e., contains more
than 3 mole % propane) then it may be pre-cooled against the HFC in
pre-cooling apparatus 101, scrubbed in a scrub column 103 (of FIG.
1) that removes excess propane and other heavier hydrocarbons, and
pre-cooled further to produce the methane make up stream 117 (of
FIG. 1). This procedure ensures that the methane make up stream
used to make the mixed refrigerant contains low enough amounts of
propane for safety.
[0050] With respect to the ethane enriched stream, predominately,
as used herein, is defined as meaning that the stream comprises at
least 90 mole % ethane. Commercial high purity ethane may contain
up to 10 mole % propane while on-site prepared ethane may have a
purity much higher than 90 mole % ethane. Thus, the minimum purity
of the ethane enriched stream is 90 mole % ethane.
[0051] With respect to the nitrogen enriched stream, predominately
as used herein, is defined as meaning that the stream comprises at
least 97 mole % nitrogen and a dew point lower than -40.degree. C.
Packaged nitrogen generator units based upon membrane separation of
air have been commonly used in marine applications to provide at
least 97 mole % nitrogen stream. Up to 99.99 mole % purity of
nitrogen may be achieved with these units in an economical fashion.
The membranes are typically operated at air feed pressures of less
than 14 bara and temperatures of less than 50.degree. C.
[0052] The use of propane, which is considered to be unfavorable
for use on the FPSO due to the possibility of formation of
flammable clouds at surface level, may be eliminated, or nearly
eliminated when using HFC's as a pre-coolant.
[0053] The main liquefaction exchanger 115 may be a wound coil
exchanger, a plate-fin exchanger, or any other exchanger typical
for cryogenic service. Vapor stream 114 may enter the main
liquefaction exchanger 115 where it is condensed and sub-cooled and
exits as LNG product stream 116 at a temperature between
-140.degree. C. to -170.degree. C. and pressure between 30-85 bara,
for example.
[0054] The condensed and sub-cooled LNG product stream 116 may be
further processed by reducing its pressure in a liquid expander
(not shown) or a flash valve (not shown) to around 1.2 bara,
forming flash gas and a liquid LNG product. The LNG product may be
subsequently sent to storage, for example.
[0055] The low pressure, warm main liquefaction refrigerant stream
130 may be sent to a sequence of inter-cooled compressors 131, 135
where the stream 130 is first compressed in compressor 131 to form
stream 132, cooled in intercooler 133 to form stream 134, further
compressed in compressor 135 to form stream 138, and then further
cooled in aftercooler 139 to emerge as a high pressure fluid stream
140. Compressors 131 and 135 are driven by driver 136. Driver 136
can be an electrical motor or a gas turbine. High pressure fluid
stream 140 may be at pressures ranging between 30-80 bara and a
temperature dictated by: (1) the coolant used in the intercooler
133 and aftercooler 139; and (2) the size of the intercooler 133
and aftercooler 139. While FIG. 1 illustrates the mixed refrigerant
compression system having one intercooler 133 and one aftercooler
139, multiple intercoolers and aftercoolers may be implemented, for
example. The coolant used in the intercooler 133 and aftercooler
139 may be air, or typically for FPSO applications, sea water, or
fresh water, which is in turn cooled by sea water, for example.
[0056] The cooled high pressure refrigerant stream 140 may be
pre-cooled using pre-cooling apparatus 101 resulting in pre-cooled
stream 141. Pre-cooled stream 141 may be separated into lighter
refrigerant stream 143 and heavier refrigerant streams 144 in
separator 142. The lighter refrigerant stream 143 may then be
condensed and sub-cooled in the main liquefaction exchanger 115 to
form stream 148, expanded in J-T valve 149 to generate cryogenic
refrigerant stream 150 having a temperature between -180.degree. C.
to -120.degree. C., before it is then vaporized in the main
liquefaction exchanger 115. The heavier refrigerant liquid stream
144 may also be sub-cooled in the main liquefaction exchanger 115
to form stream 145 where it may then be expanded in J-T valve 146
to generate cryogenic refrigerant stream 147 to also be vaporized
in the main liquefaction exchanger 115. The current process may
also include a hydraulic expander (not shown) before J-T valve 146
to improve efficiency.
[0057] The combined cryogenic refrigerant streams 147, 150 boil at
successively higher temperatures while flowing down the main
liquefaction exchanger 115 before eventually exiting the exchanger
as the vapor stream 130 at or slightly above dew point thereby
completing the refrigeration loop.
[0058] FIG. 2 illustrates the internals of pre-cooling apparatus
101 comprising one or more HFC evaporators. HFC evaporators 222,
226, 230 are used to cool the dehydrated natural gas feed stream
100 to approximately -100.degree. C. to 0.degree. C. The
evaporators may be shell-and-tube type exchangers, for example.
Shell-and-tube type exchangers are also commonly termed "kettles"
since the shell side consists of a pool of boiling HFC refrigerant.
The HFC refrigerant stream 162 (also illustrated in FIG. 1) may be
saturated or preferably sub-cooled HFC liquid at a temperature
determined by the air or sea water coolant as well as the condenser
and sub-cooler size. For example, typical sea water cooled in FPSO
applications may cool the HFC discharge from multi-stage HFC
compressor 158 to 5-20.degree. C. above sea water temperatures.
[0059] The high pressure, sub-cooled HFC stream 162 may be split up
into streams 220, 240. Stream 220 may provide the cooling to the
dehydrated natural gas feed stream 100 while stream 240 may provide
the cooling to the mixed refrigerant stream 140. Stream 220 may be
subsequently expanded in a J-T valve 290 to form stream 221 and
then sent to the boiling pool of liquid in evaporator 222. The HFC
saturated vapor stream 223 exiting evaporator 222 may be combined
with the saturated vapor stream 243 arising from evaporator 242 and
the combined stream 163 (also illustrated in FIG. 1) may then be
sent as a side-stream to the highest pressure suction nozzle of the
multi-stage HFC compressor 158 of FIG. 1. The HFC liquid in 222
which has not boiled away may be sent via stream 224 to another
pressure letdown J-T valve 292 further reducing its pressure and
temperature resulting in stream 225. The boiling liquid stream 225
may be sent to evaporator 226 where it supplies further
refrigeration to stream 201. The vapor stream 227 arising from
evaporator 226 may be combined with vapor stream 247 arising from
evaporator 246 and the combined stream 164 (also illustrated in
FIG. 1) is sent as a side-stream to the mid-pressure suction nozzle
of the multi-stage HFC compressor 158. The liquid from 226 as
stream 228 may be further flashed in a J-T valve 294 resulting in
boiling stream 229 and sent to the boiling pool of liquid in
evaporator 230 where it may be evaporated completely to form
saturated vapor stream 231. Saturated vapor stream 231 may be
combined with the vapor stream 251 arising from the evaporator 250
and the combined stream 165 (also illustrated in FIG. 1) may be
sent to the low pressure suction inlet of the multi-stage HFC
compressor 158.
[0060] Evaporator 230 acts as the condenser for the scrub column
103. The temperature of the boiling liquid in evaporator 230, the
outlet pressure of the final J-T valve 294, as well as the HFC
itself may be decided by several criteria. For example, it is
typical to limit the Higher Heating Value (HHV) of the LNG that
requires the generation of sufficient reflux into the scrub column
103 which in turn requires sufficiently low boiling temperatures in
evaporator 230. Typically, this temperature varies between
-20.degree. C. to -80.degree. C. although there may be exceptions
where this temperature may be outside of this range. For example,
refrigerants like R-410A when flashed down to a minimum allowable
pressure of around 1.25 bara provide a temperature of around
-50.degree. C. In this case, the final pre-cooling temperature of
about -47.degree. C. may be reached at the outlet of evaporator
230. HFC's can also be safely used at sub-atmospheric pressures to
achieve even lower temperatures. One reason this is possible,
unlike propane, is that air-HFC mixtures are non-flammable, so that
even with some air ingress, no flammability concerns arise.
Selection of the HFC refrigerant, therefore, can also be governed
by the final pre-cooling temperature that needs to be attained to
condense enough hydrocarbons out of the feed gas.
[0061] Similar to the pre-cooling of the dehydrated natural gas
feed stream 100 above, evaporators 242, 246, 250 may be used to
pre-cool the hydrocarbon mixed refrigerant stream 140 to yield a
pre-cooled stream 141. The pre-cooled temperature of evaporator 242
may be identical to evaporator 222, the temperature of evaporator
246 may be identical to evaporator 226, and the temperature of
evaporator 250 may be identical to evaporator 230. In general, if
there are (n) cooling stages, the boiling temperature of the HFC in
any pre-cooling stage (k) will be identical for both the feed and
hydrocarbon mixed refrigerant evaporators. In this way, the
saturated HFC vapor leaving both the feed and hydrocarbon mixed
refrigerant evaporator in any given pre-cooling stage will be at
identical pressures. This allows the vapor streams to be mixed and
sent to a single compressor stage, hence reducing the number of
compressor inlet nozzles.
[0062] Another advantage of using HFC's as the precoolant is that
HFC's have a critical temperature that exceeds 50.0.degree. C.
Because the HFC's critical temperature exceeds 50.0.degree. C., the
HFC condenser 161 (illustrated in FIG. 1) following the multi-stage
HFC compressor 158 does not operate in the supercritical region.
Because the HFC condenser 161 following the multi-stage HFC
compressor 158 does not operate in the supercritical region, the
system will, thus, operate at a higher efficiency since the
volumetric flow reduction is larger at sub-critical conditions.
Further HFC's can provide refrigeration to the feed and the main
liquefaction refrigerant in several boiling stages which allows the
system to match the cooling curve to an arbitrary degree of
closeness thereby resulting in higher efficiencies.
[0063] The HFC may be used to provide condensing duties at lower
temperatures than propane for a hydrocarbon scrub column reflux,
thereby simplifying the design of the liquefaction plant and taking
a portion of the load off the hydrocarbon mixed refrigerant. This
relief of the hydrocarbon mixed refrigerant will improve the
efficiency of the system and reduce the size and capital cost of
the expensive cryogenic section of the plant. Moreover, simpler and
more cost effective process configurations may be used over a wider
range of feed compositions and product specifications.
[0064] In another embodiment as illustrated in FIG. 3, the HFC
pre-cooling system is not used for supplying the condensing duty
for the scrub column 303. Instead, the hydrocarbon rich mixed
refrigerant is used for supplying the condensing duty for the scrub
column 303. The dehydrated natural gas feed stream 300 is
pre-cooled in two to four pre-cooling stages in pre-cooling
apparatus 301 to yield a pre-cooled stream 302. The temperature of
pre-cooled stream 302 depends upon the reflux rate in the scrub
column 303 and the level of hydrocarbon removal required. For a
given reflux rate, lowering the temperature of the pre-cooled
stream 302 increases the hydrocarbon removal rate in the column.
The overhead stream 304 exiting the scrub column 303 enters the
main liquefaction exchanger 313 where it is cooled and partially
liquefied resulting in stream 306. Stream 306 is phase-separated in
the scrub column reflux drum 307 resulting in streams 308 and 312.
Stream 308 is pumped in pump 309 and then sent via stream 310 to
the scrub column 303 as column reflux. Stream 312 is re-inserted
back into the main liquefaction exchanger 313 for liquefaction, to
become LNG product stream 316. It should be noted that the
remaining streams and apparatus shown in FIG. 3 function similar to
the equivalent 100 series streams and apparatus shown in FIG.
1.
[0065] FIGS. 4A and 4B illustrate two HFC compressor driver
configurations. FIG. 4A shows a configuration in which the HFC
compressor 158 is driven directly by a multi-shaft gas turbine
driver 154 (also illustrated in FIG. 1). Aero-derivative turbines
such as LM2500, LM6000, or PGT25 are commercially available
examples of multi-shaft turbines. The multi-shaft turbine driver
154 includes a gas generator 456 and a power turbine expander 457
on separate shafts. The power turbine 457 requires a high starting
torque since it is loaded with the multi-stage HFC compressor 158
that is started without venting any HFC inventory. At start-up, the
air compressor 454, driven by the gas generator 456, is allowed to
develop full pressure before starting power turbine 457. Once the
air compressor 454 develops full pressure, the still high pressure
exhaust gas discharging from the gas generator 456 provides the
required high starting torque to power turbine 457. In addition to
having high starting torque, multi-shaft aero-derivative turbines
also have significant speed flexibility due to the flexibility to
control gas pressure at the exit of combustor 455, at the exit of
gas generator 456, and at the exit of power turbine 457. Yet
another way to provide a high starting torque for the HFC
compressor 158 is to use a steam turbine directly coupled to the
multi-stage HFC compressor 158.
[0066] FIG. 4B illustrates an electric motor driven multi-stage HFC
compressor 158. The motor 466 is supplied power from a power supply
464 either through connectivity with the power grid on the FPSO or
through connectivity with an electrical generator driven by a
multi-shaft gas turbine. There may be a variable speed drive (VSD)
that controls the current, voltage, and frequency across the stator
coils of the motor to prevent coil damage as well as prevent
instability in the grid. For a typical motor, the starting torque
(or rotor lock-up torque) is proportional to .phi..times.I.sub.Rot
where .phi. is the magnetic flux intensity from the stator coils
and I.sub.Rot is the rotor induced current. In turn .phi. is
proportional to the stator voltage. At zero rotor speed, I.sub.Rot
is proportional to the frequency of rotation of the stator field
which in turn is determined by the frequency of the input stator
voltage and the number of stator poles (i.e., 120.times.(f/p))
where f=frequency of the input stator voltage and p=number of
stator poles. Hence, by applying a sufficiently high frequency and
voltage in the stator windings, the rotor can be imparted a
starting torque that overcomes load torque consisting of the
compressor and the HFC gas that is contained within it. Variable
Frequency Drive (VFD) technologies may be used to adjust the input
frequency to the stator coils while maintaining the
voltage-to-frequency ratio with the goal of providing sufficiently
high starting torque.
[0067] A concern with HFC compressors occurs due to the HFC's
higher molecular weights in comparison to conventional pre-cooling
refrigerants used in LNG service like propane, for example. While
propane has a molecular weight of 44.096, an HFC like R410A, which
satisfies toxicity, flammability, low temperature glide, and other
desired properties, has a molecule weight of 72.58. It is
well-known that the velocity of sound in a fluid is calculated
by:
C = ( .gamma. R T M W ) 0.5 ( 1 ) ##EQU00001##
where .gamma.=Adiabatic exponent of the HFC, R=Universal Gas
constant (8.314 J/Mol K), T=Temperature of HFC at any point within
the compressor, and MW=Molecular weight of HFC. The velocity of
sound (C) varies inversely with the square root of the molecular
weight of the fluid. Higher molecular weight refrigerants,
therefore, give rise to lower sonic velocities. Most compressor
designs limit the impeller tip mach number (Ma) to sub-sonic levels
(i.e. less than 1.0). Ma is calculated by:
Ma=.PI..times.RPM.times.Tip Diameter/C (2)
where RPM=Revolutions Per Minute of the impeller, .PI.=3.14159, and
C=Velocity of sound as calculated from the equation (1). In order
to limit Ma to sub-sonic levels (i.e. less than 1.0), either the
compressor must be run at a low enough speed (RPM) or the impeller
diameter must be low.
[0068] Running the compressor at low speeds and/or using smaller
impeller diameters, however, limits the theoretical polytrophic
pressure ratio per stage as shown by the well-established
centrifugal compressor expression:
P Out P In + { 1 + ( n - 1 ) M W n Z Avg R T In U T V t , Out } n (
n - 1 ) ( 3 ) ##EQU00002##
where n=HFC Polytrophic co-efficient, Z.sub.Avg=Average HFC
Compressibility Factor, R=Universal Gas Constant (8.314 J/mol.K),
T.sub.In=Temperature at inlet of impeller, P.sub.In=Pressure at
inlet of impeller, MW=Molecular weight of the gas, and
V.sub.t,Out=Absolute tangential velocity of HFC at the impeller
tip. The above pressure ratio assumes that the flow into the
impeller inlet passage has no tangential component. This velocity
is given by the well-known expression obtained by the method of
velocity triangles:
V t , Out = ( U T - Q tan .beta. .pi. D T b ) ( 4 )
##EQU00003##
where Q=Volumetric Flow Rate, .beta.=Impeller tip angle with the
radial direction, D.sub.T=Impeller tip diameter, and b=Impeller
average flow channel width. The theoretical power input per unit
mass required is given by:
Power per Stage=V.sub.t,Out.times.U.sub.T. (5)
[0069] To achieve higher pressure ratios required to reach the
condensing pressure of the HFC, several impeller-diffuser (wheels)
stages may be required which implies that the there may need to be
several compressor casings which adds to the capital cost of the
plant. It was determined, however, that single casing HFC
compressors can be designed for HFC's like R410A. Table 3 shows the
impact of differing HFC's and their comparison with conventional
pre-cooling refrigerants like propane on the pressure ratio of a
stage for a flow with Q=75,000 m3/h, T.sub.In=10.00.degree. C., and
P.sub.In=2.92 bara which is typical for one of the pressure levels
of an HFC pre-cooled process.
TABLE-US-00003 TABLE 3 Vt Mol C Ut Tip Out Type Component Wt. Gamma
Z (m/s) (m/s) Ma (m/s) Pout/Pin Propane C.sub.3H.sub.8 44.10 1.17
0.94 249.66 115.45 0.46 69.48 1.17 R410A F32/F125 72.37 1.21 0.96
197.95 115.45 0.58 69.48 1.29 R134a C.sub.2H.sub.2F.sub.4 102.03
1.15 0.93 163.05 115.45 0.71 69.48 1.44 R125 C.sub.2HF.sub.5 120.02
1.13 0.95 148.71 115.45 0.78 69.48 1.53
For the above analysis, the impeller angle .beta. was 20 degrees,
the tip diameter was 1.05 m and the average channel width was 0.05
m. The impeller rotated at 2100 RPM. The stage is assumed to be
approximately isentropic for simplicity of demonstration. This is
especially true for low values of .beta. which avoids separation of
the boundary layer along the length of the compressor blade flow
passage.
[0070] Table 3 illustrates that for this impeller geometry and
identical inlet flow conditions, the Mach number increases with an
increase in molecular weight. Table 3 also illustrates that the
stage pressure ratio is strongly dependent upon the molecular
weight. A compressor achieves higher pressure ratio with HFC's in
comparison to propane for the same inlet volumetric flow,
temperature, and pressure. Table 3 shows significant advantage of
some HFC's like R410A over conventional refrigerants like propane
in terms of higher pressure ratios per stage for the same
compressor geometry and input power.
EXAMPLES
[0071] In this exemplary illustration, a dehydrated natural gas
feed stream 100 was liquefied to form LNG using the HFC pre-cooled
mixed refrigerant process. The total dehydrated natural gas feed
stream 100 entering the pre-cooling section was around 2.39 mmtpa
or 301.2 tph. The dehydrated natural gas feed stream 100 from the
drier beds (not shown in FIG. 1) entered the HFC cooled evaporator
222 at 15.degree. C. and 68.95 bara where it was cooled to a
temperature of -4.06.degree. C. The feed was further cooled in
evaporator 226 to a temperature of approximately -24.39.degree. C.
after which it was sent to the scrub column 103 as stream 102. The
cold reflux stream 111 in the scrub column 103 caused the stripping
of the C3+ (propane, butanes, pentane etc.) hydrocarbons from the
input stream 102 to yield a vapor overhead stream 104 and a bottoms
heavy hydrocarbon stream 105.
[0072] Table 4 summaries the various stream conditions from the
exemplary illustration:
TABLE-US-00004 TABLE 4 Stream 102 104 105 Flow lbmol/h 38,298.94
35,757.67 4,819.31 Temperature .degree. C. -24.38 -27.56 -24.47
Pressure bara 68.31 68.27 68.29 Vapor % 89.40 100.00 0.00 Percent
Dew Point .degree. C. -27.56 Bubble Point .degree. C. -57.90 -24.47
Nitrogen Mole % 0.95 1.01 0.22 Methane Mole % 85.32 89.11 48.04
Ethane Mole % 6.81 6.33 14.02 Propane Mole % 3.73 2.66 15.61
Isobutane Mole % 0.48 0.22 2.90 n-Butane Mole % 1.63 0.59 10.73
Isopentane Mole % 0.35 0.04 2.69 n-Pentane Mole % 0.36 0.03 2.78
Hexane Mole % 0.38 0.00 3.01 Carbon Mole % 0.01 0.01 0.01
Dioxide
The heavy hydrocarbon stream 105 was typically led to a
de-methanizer column where lighter components like methane and
ethane were ejected to produce additional fuel gas while the
remaining stream rich in C3+ was either further fractionated or
stored and shipped. Stream 107 entered the evaporator 230 to cool
it to -44.39.degree. C. to generate partially condensed stream 108,
the liquid portion (stream 111) of which was used as scrub column
reflux 103, while the vapor stream 114 was sent to the main
liquefaction exchanger 115.
[0073] Table 5 contains the summary of streams 108, 111, and 114
around the scrub column condenser 110.
TABLE-US-00005 TABLE 5 Stream 108 114 111 Flow lbmol/h 30,394.02
33,479.63 2,278.04 Temperature .degree. C. -44.39 -42.30 -42.30
Pressure Bara 67.79 67.79 67.79 Vapor Percent % 91.89 100.00 0.00
Dew Point .degree. C. -42.30 Bubble Point .degree. C. -52.30 -42.30
Nitrogen Mole % 0.95 1.01 0.22 Methane Mole % 85.32 89.11 48.04
Ethane Mole % 6.81 6.33 14.02 Propane Mole % 3.73 2.66 15.61
Isobutane Mole % 0.48 0.22 2.90 n-Butane Mole % 1.63 0.59 10.73
Isopentane Mole % 0.35 0.04 2.69 n-Pentane Mole % 0.36 0.03 2.78
Hexane Mole % 0.38 0.00 3.01 Carbon Dioxide Mole % 0.01 0.01
0.01
Table 5 shows that the vapor stream 114, at its dew point, was sent
to the main liquefaction exchanger 115 where it was fully condensed
and sub-cooled against hydrocarbon rich mixed refrigerant streams
in order to yield stream 116 at a temperature of -152.72.degree. C.
Stream 116 was then expanded in an LNG hydraulic expander (details
not shown) to extract some power as well as further expanded it in
a J-T valve to around 1.17-1.38 bara generating flash gas and a
final LNG product.
[0074] The composition of the hydrocarbon rich mixed refrigerant
(HMR) stream 130 was 16.6 mole % N.sub.2, 40.99 mole % CH.sub.4,
and 42.4 mole % C.sub.2H.sub.6, with a total flow of 65,758
lbmol/h. This mixture had a dew point of -62.28.degree. C. and a
bubble point of -157.87.degree. C. Warm HMR stream 130 (where
T=-59.5.degree. C., P=8.47 bara) was led to the suction of the low
pressure compressor 131. Two inter-cooled compression stages
boosted its pressure to yield stream 138 (T=46.50.degree. C.,
P=61.33 bara) with a dew point of -10.33.degree. C. and a bubble
point of -76.89.degree. C. The HMR stream, therefore, was further
cooled to -76.89.degree. C. to convert it into a saturated liquid
and further sub-cooled to provide the required cryogenic
refrigeration temperatures. This cooling was done using
after-cooler 139, HFC in evaporators 242, 246, 250, and the boiling
HMR in the main liquefaction exchanger 115. HMR stream 141 exited
the lowest temperature HFC evaporator 250 at -44.39.degree. C. and
59.23 bara. The partly condensed stream 141 was phase-separated in
drum 142 to yield a lighter hydrocarbon HMR stream 143 and a
heavier HMR stream 144. Stream 143 was then condensed and
sub-cooled in the main liquefaction exchanger 115 to yield the cold
refrigerant stream 148 while stream 144 was sub-cooled to yield the
refrigerant stream 145. Streams 145, 148 were further expanded in
J-T valves 146, 149 to yield the low pressure refrigerant streams
147, 150 that are then re-introduced into the main liquefaction
exchanger 115 to provide refrigeration to the feed stream 114 and
HMR streams 143, 144.
[0075] Table 6 shows the stream conditions of various HMR streams
in this circuit.
TABLE-US-00006 TABLE 6 Stream 130 140 141 143 144 Flow lbmol/h
65,758.04 65,758.04 65,758.04 24,816.49 40,941.55 Temperature
.degree. C. -59.50 15.00 -44.42 -44.50 -44.50 Pressure bara 8.47
60.73 59.23 59.03 59.03 Vapor Percent % 100.00 100.00 37.62 100.00
0.00 Dew Point .degree. C. -62.28 -10.53 -11.04 -44.50 3.46 Bubble
Point .degree. C. -157.87 -77.65 -79.60 -99.04 -44.50 Nitrogen Mole
% 16.58 16.58 16.58 30.92 7.89 Methane Mole % 40.99 40.99 40.99
51.12 34.85 Ethane Mole % 42.43 42.43 42.43 17.95 57.26
[0076] The HFC pre-cooling loop is now described. Referring to FIG.
2, there were 3 stages of evaporation. The first stage included
evaporators 222, 242 that cool the dehydrated natural gas feed
stream 100 and HMR stream 140 to -4.06.degree. C. by boiling R410A
streams 221, 241 supplied at -7.11.degree. C. and 6.38 bara. The
vapor streams 223, 243 from the evaporators 222, 242 respectively
are sent to the high pressure suction of R410A compressor 158. The
un-evaporated liquid streams 224, 244 were isenthalpically flashed
down to 3.03 bara and -27.39.degree. C. resulting in respective
streams 225, 245. These streams were sent respectively to
evaporators 226, 246 to cool the feed and mixed refrigerant down to
-24.39.degree. C. The vapor streams 227, 247 leaving evaporators
226, 246 were led to the mid-pressure suction nozzle of R410A
compressor 158. The liquid streams 228, 248 leaving evaporators
226, 246 were isenthalpically flashed down to 1.25 bara and
-47.39.degree. C. resulting in low pressure refrigerant streams
229, 249 respectively. Streams 229, 249 were completely vaporized
to streams 231, 251 respectively and sent to the low pressure inlet
nozzle of R410A compressor 158. The R410A compressor 158 was a
3-stage machine which compressed the low 165, mid-level 164, and
high 163 pressure vapor streams to a final pressure of 15.58 bara
and 65.11.degree. C. in stream 160. Stream 160 was then
de-superheated, condensed, and sub-cooled against sea water in
exchanger 161 to yield the condensed R410A stream 162 at 13.58 bara
and 15.degree. C. and re-circulated to the R410A evaporators
completing the refrigeration loop. In this case the total
circulation rate of the R410A was 46,119 lbmol/h.
[0077] The total LNG production was about 2.39 million metric tones
per annum (mmtpa). The R410A compressor 158 required 27 MW. The
inter-stage cooled MR compressors 133 and 135 required 40.5 MW. A
specific power of approximately 270.26 kwh/ton was realized.
[0078] Various turbines suitable for the marine environment or
electric motors may be used as drivers. For instance, a GE LM6000
at 26 C yielding de-rated power output of about 28 MW can be used
as a driver for the R410A compressor 158 as well as 2.times.50% MR
compressor strings. Gear boxes may be utilized to adjust the speed
of the R410A compressor which tends to run at a lower speeds due to
impeller tip Mach number restrictions.
[0079] A hydrofluorocarbon within a pure fluid cascade process
consisting of a plurality of refrigeration loops similar to the
ConocoPhillips cascade process outlined in U.S. Pat. No. 5,669,234,
and incorporated herein by reference, may also be used. This
process involves three cooling loops: (1) a pre-cooling loop using
propane which pre-cools the feed and a lower boiling refrigerant
like ethane or ethylene; (2) an intermediate cooling loop using
ethane as refrigerant that cools the feed further and a lower
boiling refrigerant like methane; and (3) a sub-cooling loop that
uses methane as refrigerant and sub-cools the feed to LNG
temperatures. In one embodiment, the propane in the pre-cooling
loop of the cascade cycle may be replaced with one of the HFC
refrigerants. The ethane in the cascade loop with a normal boiling
point of -89.degree. C. may be replaced by a lower boiling HFC like
R23 (CHF.sub.3) with a normal boiling point of -82.1.degree. C.
[0080] The process of FIG. 1 was also used to compare the energy
efficiencies in terms of kwh/ton LNG, obtained from three different
pre-cooling refrigerants: R410A, propane, and CO.sub.2 with and
without the use of propane in the mixed refrigerant. The feed flow
rate was maintained fixed and the total power consumption of the
pre-cooling and mixed refrigerant compressors was minimized. For
uniformity of comparisons, the pre-cooling temperatures were set to
at least -37.degree. C. in order to condense sufficient amounts of
hydrocarbons from the dehydrated natural gas feed stream 100. In
some cases, the optimal pre-cooling temperatures were found to be
lower than -38.degree. C. Table 7 shows these results.
[0081] The first three processes included propane in the mixed
refrigerant. The next three processes illustrate having no propane
in the mixed refrigerant. For the final three processes, it was
assumed from Table 5 (stream 114) that the methane stream used to
prepare the mixed refrigerant contains 2.66 mole % propane. It was
also assumed that an on-site de-ethanizer column may be used for
producing an ethane stream that contains 1.0% propane and 99.0%
ethane, which is typical in the industry. A typical mixed
refrigerant with a composition of 40.99% methane and 42.4% ethane
will result in a propane content of
0.4099.times.2.66+0.424.times.1.00=1.51 mole %.
TABLE-US-00007 TABLE 7 ASHRAE Specific Temperature Pressure
Toxicity & Power Mole % of Streams of Streams Flammability
Pre-cooling Liquefaction (kwh/ton Propane 229 and 249 229 and Index
Refrigerant Refrigerant LNG) in MR (C.) 249 (bara) A1 CO.sub.2 MR
(with 262.26 11.80 -37.09 11.17 propane) A3 Propane MR (with 250.89
12.42 -37.36 1.24 propane) A1 R410A MR (with 249.09 9.93 -37.29
1.99 propane) A1 CO.sub.2 MR (without 270.91 0.00 -37.22 11.12
propane) A3 Propane MR (without 256.59 0.00 -37.36 1.24 propane) A1
R410A MR (without 251.96 0.00 -38.55 1.82 propane) A1 CO2 MR (with
273.76 1.51 -37.76 10.90 limited propane) A3 Propane MR (with
254.65 1.51 -36.77 1.27 limited propane) A1 R410A MR (with 250.44
1.51 -39.01 1.78 limited propane)
[0082] First, the processes using CO.sub.2 had a specific power
consumption approximately 5-9% higher than the ones not using
CO.sub.2, rendering the former less efficient than processes based
upon R410A and propane pre-cooling. Second, the specific power of
the process using R410A mixed refrigerant (excluding propane) was
about 0.42% higher than that of the process using C3MR (including
propane). Process C3MR (including propane) was used for large
land-based plants with the highest possible energy efficiency. This
comparison demonstrates that the power penalty incurred in
eliminating propane entirely from the liquefaction cycle was
insignificant and that the energy efficiency of the R410A based
process may be considered to be comparable to the most efficient
land-based plants. Third, the reduction of propane inventory was
significant since the optimized mixed refrigerant composition for
the processes using propane in the mixed refrigerant involves 9-12
mol % propane. Propane inventory and capital costs are also further
reduced by eliminating the distillation column and storage
typically provided to make propane.
[0083] Embodiments of the current invention could also be applied
to replacing the hydrocarbon mixed refrigerant mixture with a
mixture of HFC components and hydrocarbons and an inert gas like
N.sub.2. In such a process, pre-cooling would be performed by a
constant boiling refrigerant (such as a single component HFC or an
azeotropic HFC) while the liquefaction and sub-cooling of the feed
would be performed using a mixture of HFC components and nitrogen.
The propane in the hydrocarbon mixed refrigerant mixture could be
replaced by an HFC like R410A or R134A, for example. The ethane
could be replaced by R23 and the methane could be replaced by R14.
In such a process, the HFC mixed refrigerant would be compressed
and cooled with air or sea water resulting in partially condensed
refrigerant which would be separated to yield a nitrogen and low
boiling point rich vapor (stream 143 in FIGS. 1 and 3) and a high
boiling HFC rich liquid (144 in FIGS. 1 and 3). The vapor and the
liquid would then be further cooled and then expanded to provide
refrigeration to the feed.
[0084] It should also be noted that because the HFC refrigerant may
be used to replace the propane refrigerant in the traditional C3MR
process, the vast experience with the design and operation of the
C3MR process can, therefore, be still exploited. For example, the
HFC pre-cooling refrigerant loop consists of several evaporators in
series whose operation is well-established in the traditional C3MR
system and process.
[0085] Finally, it has been verified that single casing compressors
with side-streams can be designed for HFC service running at speeds
between 1700-5000 RPMs which is the normal operating range for a
vast array of commercially available low heat rate aero-derivative
turbines.
[0086] While aspects of the present invention has been described in
connection with the preferred embodiments of the various figures,
it is to be understood that other similar embodiments may be used
or modifications and additions may be made to the described
embodiment for performing the same function of the present
invention without deviating therefrom. Therefore, the claimed
invention should not be limited to any single embodiment, but
rather should be construed in breadth and scope in accordance with
the appended claims.
* * * * *