U.S. patent application number 12/423306 was filed with the patent office on 2009-11-19 for liquefied natural gas and hydrocarbon gas processing.
This patent application is currently assigned to ORTLOFF ENGINEERS, LTD.. Invention is credited to Kyle T. Cuellar, Hank M. Hudson, Tony L. Martinez, John D. Wilkinson.
Application Number | 20090282865 12/423306 |
Document ID | / |
Family ID | 41314848 |
Filed Date | 2009-11-19 |
United States Patent
Application |
20090282865 |
Kind Code |
A1 |
Martinez; Tony L. ; et
al. |
November 19, 2009 |
Liquefied Natural Gas and Hydrocarbon Gas Processing
Abstract
A process for the recovery of ethane, ethylene, propane,
propylene, and heavier hydrocarbons from a liquefied natural gas
(LNG) stream and a hydrocarbon gas stream is disclosed. The LNG
feed stream is divided into two portions. The first portion is
supplied to a fractionation column at a first upper mid-column feed
point. The second portion is directed in heat exchange relation
with a first portion of a warmer distillation stream rising from
the fractionation stages of the column, whereby the LNG feed stream
is partially heated and the distillation stream is totally
condensed. The condensed distillation stream is divided into a
"lean" LNG stream and a reflux stream, whereupon the reflux stream
is supplied to the column at a top column feed position. The second
portion of the LNG feed stream is heated further to partially or
totally vaporize it and thereafter supplied to the column at a
first lower mid-column feed position. The gas stream is divided
into two portions. The second portion is expanded to the operating
pressure of the column, then both portions are directed in heat
exchange relation with the lean LNG stream and the second portion
of the warmer distillation stream, whereby both portions of the gas
stream are cooled, the lean LNG stream is vaporized, and the second
portion of the distillation stream is heated. The first portion of
the gas stream, which has been cooled to substantial condensation,
is supplied to the column at a second upper mid-column feed point,
and the second portion is supplied to the column at a second lower
mid-column feed point. The quantities and temperatures of the feeds
to the column are effective to maintain the column overhead
temperature at a temperature whereby the major portion of the
desired components is recovered in the bottom liquid product from
the column.
Inventors: |
Martinez; Tony L.; (Odessa,
TX) ; Wilkinson; John D.; (Midland, TX) ;
Hudson; Hank M.; (Midland, TX) ; Cuellar; Kyle
T.; (Katy, TX) |
Correspondence
Address: |
FITZPATRICK CELLA HARPER & SCINTO
1290 Avenue of the Americas
NEW YORK
NY
10104-3800
US
|
Assignee: |
ORTLOFF ENGINEERS, LTD.
Midland
TX
|
Family ID: |
41314848 |
Appl. No.: |
12/423306 |
Filed: |
April 14, 2009 |
Related U.S. Patent Documents
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Application
Number |
Filing Date |
Patent Number |
|
|
61053814 |
May 16, 2008 |
|
|
|
Current U.S.
Class: |
62/620 ;
62/634 |
Current CPC
Class: |
F25J 2290/50 20130101;
F25J 2270/904 20130101; F25J 2240/02 20130101; F25J 2210/62
20130101; F25J 2205/04 20130101; F25J 2200/72 20130101; F25J
2230/08 20130101; F25J 3/0214 20130101; F25J 3/0209 20130101; F25J
2200/78 20130101; F25J 2200/02 20130101; F25J 3/0238 20130101; F25J
2200/76 20130101; F25J 3/0233 20130101; F25J 2210/02 20130101; F25J
2230/60 20130101; F25J 2200/38 20130101; F25J 2210/06 20130101;
F25J 2235/60 20130101 |
Class at
Publication: |
62/620 ;
62/634 |
International
Class: |
F25J 3/00 20060101
F25J003/00 |
Claims
1. A process for the separation of liquefied natural gas containing
methane and heavier hydrocarbon components and a gas stream
containing methane and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said
methane and a relatively less volatile liquid fraction containing a
major portion of said heavier hydrocarbon components wherein (a)
said liquefied natural gas is divided into at least a first liquid
stream and a second liquid stream; (b) said first liquid stream is
expanded to lower pressure and is thereafter supplied to a
distillation column at a first upper mid-column feed position; (c)
said second liquid stream is heated sufficiently to vaporize it,
thereby forming a vapor stream; (d) said vapor stream is expanded
to said lower pressure and is supplied to said distillation column
at a first lower mid-column feed position; (e) said gas stream is
divided into at least a first gaseous stream and a second gaseous
stream; (f) said first gaseous stream is cooled to condense
substantially all of it and is thereafter expanded to said lower
pressure whereby it is further cooled; (g) said expanded
substantially condensed first gaseous stream is thereafter supplied
to said distillation column at a second upper mid-column feed
position; (h) said second gaseous stream is expanded to said lower
pressure, is cooled, and is thereafter supplied to said
distillation column at a second lower mid-column feed position; (i)
an overhead distillation stream is withdrawn from an upper region
of said distillation column and divided into at least a first
portion and a second portion, whereupon said first portion is
compressed to higher pressure; (j) said compressed first portion is
cooled sufficiently to at least partially condense it and form
thereby a condensed stream, with said cooling supplying at least a
portion of said heating of said second liquid stream; (k) said
condensed stream is divided into at least a volatile liquid stream
and a reflux stream; (l) said reflux stream is further cooled, with
said cooling supplying at least a portion of said heating of said
second liquid stream; (m) said further cooled reflux stream is
supplied to said distillation column at a top column feed position;
(n) said volatile liquid stream is heated sufficiently to vaporize
it, with said heating supplying at least a portion of said cooling
of one or more of said first gaseous stream and said expanded
second gaseous stream; (o) said second portion is heated, with said
heating supplying at least a portion of said cooling of one or more
of said first gaseous stream and said expanded second gaseous
stream; (p) said vaporized volatile liquid stream and said heated
second portion are combined to form said volatile residue gas
fraction containing a major portion of said methane; and (q) the
quantity and temperature of said reflux stream and the temperatures
of said feeds to said distillation column are effective to maintain
the overhead temperature of said distillation column at a
temperature whereby the major portion of said heavier hydrocarbon
components is recovered in said relatively less volatile liquid
fraction by fractionation in said distillation column.
2. A process for the separation of liquefied natural gas containing
methane and heavier hydrocarbon components and a gas stream
containing methane and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said
methane and a relatively less volatile liquid fraction containing a
major portion of said heavier hydrocarbon components wherein (a)
said liquefied natural gas is divided into at least a first liquid
stream and a second liquid stream; (b) said first liquid stream is
expanded to lower pressure and is thereafter supplied to a
distillation column at a first upper mid-column feed position; (c)
said second liquid stream is heated sufficiently to vaporize it,
thereby forming a first vapor stream; (d) said first vapor stream
is expanded to said lower pressure and thereafter supplied to said
distillation column at a first lower mid-column feed position; (e)
said gas stream is divided into at least a first gaseous stream and
a second gaseous stream; (f) said first gaseous stream is cooled to
condense substantially all of it and is thereafter expanded to said
lower pressure whereby it is further cooled; (g) said expanded
substantially condensed first gaseous stream is thereafter supplied
to said distillation column at a second upper mid-column feed
position; (h) said second gaseous stream is expanded to said lower
pressure and is thereafter cooled sufficiently to partially
condense it; (i) said partially condensed expanded second gaseous
stream is separated thereby to provide a second vapor stream and a
third liquid stream; (j) said second vapor stream is further cooled
and thereafter supplied to said distillation column at a second
lower mid-column feed position; (k) said third liquid stream is
supplied to said distillation column at a third lower mid-column
feed position; (l) an overhead distillation stream is withdrawn
from an upper region of said distillation column and divided into
at least a first portion and a second portion, whereupon said first
portion is compressed to higher pressure; (m) said compressed first
portion is cooled sufficiently to at least partially condense it
and form thereby a condensed stream, with said cooling supplying at
least a portion of said heating of said second liquid stream; (n)
said condensed stream is divided into at least a volatile liquid
stream and a reflux stream; (o) said reflux stream is further
cooled, with said cooling supplying at least a portion of said
heating of said second liquid stream; (p) said further cooled
reflux stream is supplied to said distillation column at a top
column feed position; (q) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of one or more of said first gaseous
stream, said expanded second gaseous stream, and said second vapor
stream; (r) said second portion is heated, with said heating
supplying at least a portion of said cooling of one or more of said
first gaseous stream, said expanded second gaseous stream, and said
second vapor stream; (s) said vaporized volatile liquid stream and
said heated second portion are combined to form said volatile
residue gas fraction containing a major portion of said methane;
and (t) the quantity and temperature of said reflux stream and the
temperatures of said feeds to said distillation column are
effective to maintain the overhead temperature of said distillation
column at a temperature whereby the major portion of said heavier
hydrocarbon components is recovered in said relatively less
volatile liquid fraction by fractionation in said distillation
column.
3. A process for the separation of liquefied natural gas containing
methane and heavier hydrocarbon components and a gas stream
containing methane and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said
methane and a relatively less volatile liquid fraction containing a
major portion of said heavier hydrocarbon components wherein (a)
said liquefied natural gas is divided into at least a first liquid
stream and a second liquid stream; (b) said first liquid stream is
expanded to lower pressure and is thereafter supplied to a
distillation column at a first upper mid-column feed position; (c)
said second liquid stream is heated sufficiently to partially
vaporize it; (d) said partially vaporized second liquid stream is
separated thereby to provide a vapor stream and a third liquid
stream; (e) said vapor stream is expanded to said lower pressure
and is supplied to said distillation column at a first lower
mid-column feed position; (f) said gas stream is divided into at
least a first gaseous stream and a second gaseous stream; (g) said
first gaseous stream is cooled to condense substantially all of it
and is thereafter expanded to said lower pressure whereby it is
further cooled; (h) said expanded substantially condensed first
gaseous stream is thereafter supplied to said distillation column
at a second upper mid-column feed position; (i) said second gaseous
stream is expanded to said lower pressure, is cooled, and is
thereafter supplied to said distillation column at a second lower
mid-column feed position; (j) said third liquid stream is expanded
to said lower pressure and thereafter supplied to said distillation
column at a third lower mid-column feed position; (k) an overhead
distillation stream is withdrawn from an upper region of said
distillation column and divided into at least a first portion and a
second portion, whereupon said first portion is compressed to
higher pressure; (l) said compressed first portion is cooled
sufficiently to at least partially condense it and form thereby a
condensed stream, with said cooling supplying at least a portion of
said heating of said second liquid stream; (m) said condensed
stream is divided into at least a volatile liquid stream and a
reflux stream; (n) said reflux stream is further cooled, with said
cooling supplying at least a portion of said heating of said second
liquid stream; (o) said further cooled reflux stream is supplied to
said distillation column at a top column feed position; (p) said
volatile liquid stream is heated sufficiently to vaporize it, with
said heating supplying at least a portion of said cooling of one or
more of said first gaseous stream and said expanded second gaseous
stream; (q) said second portion is heated, with said heating
supplying at least a portion of said cooling of one or more of said
first gaseous stream and said expanded second gaseous stream; (r)
said vaporized volatile liquid stream and said heated second
portion are combined to form said volatile residue gas fraction
containing a major portion of said methane; and (s) the quantity
and temperature of said reflux stream and the temperatures of said
feeds to said distillation column are effective to maintain the
overhead temperature of said distillation column at a temperature
whereby the major portion of said heavier hydrocarbon components is
recovered in said relatively less volatile liquid fraction by
fractionation in said distillation column.
4. A process for the separation of liquefied natural gas containing
methane and heavier hydrocarbon components and a gas stream
containing methane and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said
methane and a relatively less volatile liquid fraction containing a
major portion of said heavier hydrocarbon components wherein (a)
said liquefied natural gas is divided into at least a first liquid
stream and a second liquid stream; (b) said first liquid stream is
expanded to lower pressure and is thereafter supplied to a
distillation column at a first upper mid-column feed position; (c)
said second liquid stream is heated sufficiently to partially
vaporize it; (d) said partially vaporized second liquid stream is
separated thereby to provide a first vapor stream and a third
liquid stream; (e) said first vapor stream is expanded to said
lower pressure and thereafter supplied to said distillation column
at a first lower mid-column feed position; (f) said gas stream is
divided into at least a first gaseous stream and a second gaseous
stream; (g) said first gaseous stream is cooled to condense
substantially all of it and is thereafter expanded to said lower
pressure whereby it is further cooled; (h) said expanded
substantially condensed first gaseous stream is thereafter supplied
to said distillation column at a second upper mid-column feed
position; (i) said second gaseous stream is expanded to said lower
pressure; (j) said expanded second gaseous stream is cooled
sufficiently to partially condense it; (k) said partially condensed
expanded second gaseous stream is separated thereby to provide a
second vapor stream and a fourth liquid stream; (l) said second
vapor stream is further cooled and thereafter supplied to said
distillation column at a second lower mid-column feed position; (m)
said third liquid stream is expanded to said lower pressure and
thereafter supplied to said distillation column at a third lower
mid-column feed position; (n) said fourth liquid stream is supplied
to said distillation column at a fourth lower mid-column feed
position; (o) an overhead distillation stream is withdrawn from an
upper region of said distillation column and divided into at least
a first portion and a second portion, whereupon said first portion
is compressed to higher pressure; (p) said compressed first portion
is cooled sufficiently to at least partially condense it and form
thereby a condensed stream, with said cooling supplying at least a
portion of said heating of said second liquid stream; (q) said
condensed stream is divided into at least a volatile liquid stream
and a reflux stream; (r) said reflux stream is further cooled, with
said cooling supplying at least a portion of said heating of said
second liquid stream; (s) said further cooled reflux stream is
supplied to said distillation column at a top column feed position;
(t) said volatile liquid stream is heated sufficiently to vaporize
it, with said heating supplying at least a portion of said cooling
of one or more of said first gaseous stream, said expanded second
gaseous stream, and said second vapor stream; (u) said second
portion is heated, with said heating supplying at least a portion
of said cooling of one or more of said first gaseous stream, said
expanded second gaseous stream, and said second vapor stream; (v)
said vaporized volatile liquid stream and said heated second
portion are combined to form said volatile residue gas fraction
containing a major portion of said methane; and (w) the quantity
and temperature of said reflux stream and the temperatures of said
feeds to said distillation column are effective to maintain the
overhead temperature of said distillation column at a temperature
whereby the major portion of said heavier hydrocarbon components is
recovered in said relatively less volatile liquid fraction by
fractionation in said distillation column.
5. The process according to claim 1 or 3 wherein (a) said second
portion is compressed to higher pressure; (b) said compressed
second portion is heated, with said heating supplying at least a
portion of said cooling of one or more of said first gaseous stream
and said expanded second gaseous stream; and (c) said vaporized
volatile liquid stream and said heated compressed second portion
are combined to form said volatile residue gas fraction.
6. The process according to claim 2 or 4 wherein (a) said second
portion is compressed to higher pressure; (b) said compressed
second portion is heated, with said heating supplying at least a
portion of said cooling of one or more of said first gaseous
stream, said expanded second gaseous stream, and said second vapor
stream; and (c) said vaporized volatile liquid stream and said
heated compressed second portion are combined to form said volatile
residue gas fraction.
7. The process according to claim 1 or 3 wherein (a) said second
gaseous stream is cooled prior to said expansion; (b) said second
portion is compressed to higher pressure; (c) said volatile liquid
stream is heated sufficiently to vaporize it, with said heating
supplying at least a portion of said cooling of one or more of said
first gaseous stream, said second gaseous stream, and said expanded
second gaseous stream; (d) said compressed second portion is
heated, with said heating supplying at least a portion of said
cooling of one or more of said first gaseous stream, said second
gaseous stream, and said expanded second gaseous stream; and (e)
said vaporized volatile liquid stream and said heated compressed
second portion are combined to form said volatile residue gas
fraction.
8. The process according to claim 2 wherein (a) said second gaseous
stream is cooled sufficiently to partially condense it; (b) said
partially condensed second gaseous stream is separated thereby to
provide said second vapor stream and said third liquid stream; (c)
said second vapor stream is expanded to said lower pressure, is
cooled, and is thereafter supplied to said distillation column at
said second lower mid-column feed position; (d) said third liquid
stream is expanded to said lower pressure and thereafter supplied
to said distillation column at said third lower mid-column feed
position; (e) said second portion is compressed to higher pressure;
(f) said volatile liquid stream is heated sufficiently to vaporize
it, with said heating supplying at least a portion of said cooling
of one or more of said first gaseous stream, said second gaseous
stream, and said expanded second vapor stream; (g) said compressed
second portion is heated, with said heating supplying at least a
portion of said cooling of one or more of said first gaseous
stream, said second gaseous stream, and said expanded second vapor
stream; and (h) said vaporized volatile liquid stream and said
heated compressed second portion are combined to form said volatile
residue gas fraction.
9. The process according to claim 4 wherein (a) said second gaseous
stream is cooled sufficiently to partially condense it; (b) said
partially condensed second gaseous stream is separated thereby to
provide said second vapor stream and said fourth liquid stream; (c)
said second vapor stream is expanded to said lower pressure, is
cooled, and is thereafter supplied to said distillation column at
said second lower mid-column feed position; (d) said fourth liquid
stream is expanded to said lower pressure and thereafter supplied
to said distillation column at said fourth lower mid-column feed
position; (e) said second portion is compressed to higher pressure;
(f) said volatile liquid stream is heated sufficiently to vaporize
it, with said heating supplying at least a portion of said cooling
of one or more of said first gaseous stream, said second gaseous
stream, and said expanded second vapor stream; (g) said compressed
second portion is heated, with said heating supplying at least a
portion of said cooling of one or more of said first gaseous
stream, said second gaseous stream, and said expanded second vapor
stream; and (h) said vaporized volatile liquid stream and said
heated compressed second portion are combined to form said volatile
residue gas fraction.
10. The process according to claim 8 wherein (a) said gas stream is
cooled sufficiently to partially condense it; (b) said partially
condensed gas stream is separated thereby to provide said second
vapor stream and said third liquid stream; (c) said second vapor
stream is divided into at least said first gaseous stream and said
second gaseous stream; (d) said second gaseous stream is expanded
to said lower pressure, is cooled, and is thereafter supplied to
said distillation column at said second lower mid-column feed
position; (e) said volatile liquid stream is heated sufficiently to
vaporize it, with said heating supplying at least a portion of said
cooling of one or more of said gas stream, said first gaseous
stream, and said expanded second gaseous stream; and (f) said
compressed second portion is heated, with said heating supplying at
least a portion of said cooling of one or more of said gas stream,
said first gaseous stream, and said expanded second gaseous
stream.
11. The process according to claim 9 wherein (a) said gas stream is
cooled sufficiently to partially condense it; (b) said partially
condensed gas stream is separated thereby to provide said second
vapor stream and said fourth liquid stream; (c) said second vapor
stream is divided into at least said first gaseous stream and said
second gaseous stream; (d) said second gaseous stream is expanded
to said lower pressure, is cooled, and is thereafter supplied to
said distillation column at said second lower mid-column feed
position; (e) said volatile liquid stream is heated sufficiently to
vaporize it, with said heating supplying at least a portion of said
cooling of one or more of said gas stream, said first gaseous
stream, and said expanded second gaseous stream; and (f) said
compressed second portion is heated, with said heating supplying at
least a portion of said cooling of one or more of said gas stream,
said first gaseous stream, and said expanded second gaseous
stream.
12. A process for the separation of liquefied natural gas
containing methane and heavier hydrocarbon components and a gas
stream containing methane and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said
methane and a relatively less volatile liquid fraction containing a
major portion of said heavier hydrocarbon components wherein (a)
said liquefied natural gas is divided into at least a first liquid
stream and a second liquid stream; (b) said first liquid stream is
expanded to a first lower pressure and is thereafter supplied to a
first distillation column at an upper mid-column feed position; (c)
said second liquid stream is heated sufficiently to vaporize it,
thereby forming a vapor stream; (d) said vapor stream is expanded
to said first lower pressure and is supplied to said first
distillation column at a lower mid-column feed position; (e) a
first overhead distillation stream is withdrawn from an upper
region of said first distillation column and compressed to higher
pressure; (f) said compressed first overhead distillation stream is
cooled sufficiently to at least partially condense it and form
thereby a condensed stream, with said cooling supplying at least a
portion of said heating of said second liquid stream; (g) said
condensed stream is divided into at least a volatile liquid stream
and a reflux liquid stream; (h) said reflux liquid stream is
further cooled, with said cooling supplying at least a portion of
said heating of said second liquid stream; (i) said further cooled
reflux liquid stream is divided into at least a first reflux stream
and a second reflux stream; (j) said first reflux stream is
supplied to said first distillation column at a top column feed
position; (k) said gas stream is divided into at least a first
gaseous stream and a second gaseous stream; (l) said first gaseous
stream is cooled to condense substantially all of it and is
thereafter expanded to a second lower pressure whereby it is
further cooled; (m) said expanded substantially condensed first
gaseous stream is thereafter supplied to a second distillation
column at an upper mid-column feed position; (n) said second
gaseous stream is expanded to said second lower pressure, is
cooled, and is thereafter supplied to said second distillation
column at a lower mid-column feed position; (o) said second reflux
stream is supplied to said second distillation column at a top
column feed position; (p) a second overhead distillation stream is
withdrawn from an upper region of said second distillation column;
(q) said volatile liquid stream is heated sufficiently to vaporize
it, with said heating supplying at least a portion of said cooling
of one or more of said first gaseous stream and said expanded
second gaseous stream; (r) said second overhead distillation stream
is heated, with said heating supplying at least a portion of said
cooling of one or more of said first gaseous stream and said
expanded second gaseous stream; (s) said vaporized volatile liquid
stream and said heated second overhead distillation stream are
combined to form said volatile residue gas fraction containing a
major portion of said methane; (t) a first bottom liquid from said
first distillation column and a second bottom liquid from said
second distillation column are combined to form said relatively
less volatile liquid fraction; and (u) the quantities and
temperatures of said first and second reflux streams and the
temperatures of said feeds to said first and second distillation
columns are effective to maintain the overhead temperatures of said
first and second distillation columns at temperatures whereby the
major portion of said heavier hydrocarbon components is recovered
in said relatively less volatile liquid fraction by fractionation
in said first and second distillation columns.
13. A process for the separation of liquefied natural gas
containing methane and heavier hydrocarbon components and a gas
stream containing methane and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said
methane and a relatively less volatile liquid fraction containing a
major portion of said heavier hydrocarbon components wherein (a)
said liquefied natural gas is divided into at least a first liquid
stream and a second liquid stream; (b) said first liquid stream is
expanded to a first lower pressure and is thereafter supplied to a
first distillation column at an upper mid-column feed position; (c)
said second liquid stream is heated sufficiently to vaporize it,
thereby forming a first vapor stream; (d) said first vapor stream
is expanded to said first lower pressure and thereafter supplied to
said first distillation column at a lower mid-column feed position;
(e) a first overhead distillation stream is withdrawn from an upper
region of said first distillation column and compressed to higher
pressure; (f) said compressed first overhead distillation stream is
cooled sufficiently to at least partially condense it and form
thereby a condensed stream, with said cooling supplying at least a
portion of said heating of said second liquid stream; (g) said
condensed stream is divided into at least a volatile liquid stream
and a reflux liquid stream; (h) said reflux liquid stream is
further cooled, with said cooling supplying at least a portion of
said heating of said second liquid stream; (i) said further cooled
reflux liquid stream is divided into at least a first reflux stream
and a second reflux stream; (j) said first reflux stream is
supplied to said first distillation column at a top column feed
position; (k) said gas stream is divided into at least a first
gaseous stream and a second gaseous stream; (l) said first gaseous
stream is cooled to condense substantially all of it and is
thereafter expanded to a second lower pressure whereby it is
further cooled; (m) said expanded substantially condensed first
gaseous stream is thereafter supplied to a second distillation
column at an upper mid-column feed position; (n) said second
gaseous stream is expanded to said second lower pressure and is
thereafter cooled sufficiently to partially condense it; (o) said
partially condensed expanded second gaseous stream is separated
thereby to provide a second vapor stream and a third liquid stream;
(p) said second vapor stream is further cooled and thereafter
supplied to said second distillation column at a first lower
mid-column feed position; (q) said third liquid stream is supplied
to said second distillation column at a second lower mid-column
feed position; (r) said second reflux stream is supplied to said
second distillation column at a top column feed position; (s) a
second overhead distillation stream is withdrawn from an upper
region of said second distillation column; (t) said volatile liquid
stream is heated sufficiently to vaporize it, with said heating
supplying at least a portion of said cooling of one or more of said
first gaseous stream, said expanded second gaseous stream, and said
second vapor stream; (u) said second overhead distillation stream
is heated, with said heating supplying at least a portion of said
cooling of one or more of said first gaseous stream, said expanded
second gaseous stream, and said second vapor stream; (v) said
vaporized volatile liquid stream and said heated second overhead
distillation stream are combined to form said volatile residue gas
fraction containing a major portion of said methane; (w) a first
bottom liquid from said first distillation column and a second
bottom liquid from said second distillation column are combined to
form said relatively less volatile liquid fraction; and (x) the
quantities and temperatures of said first and second reflux streams
and the temperatures of said feeds to said first and second
distillation columns are effective to maintain the overhead
temperatures of said first and second distillation columns at
temperatures whereby the major portion of said heavier hydrocarbon
components is recovered in said relatively less volatile liquid
fraction by fractionation in said first and second distillation
columns.
14. A process for the separation of liquefied natural gas
containing methane and heavier hydrocarbon components and a gas
stream containing methane and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said
methane and a relatively less volatile liquid fraction containing a
major portion of said heavier hydrocarbon components wherein (a)
said liquefied natural gas is divided into at least a first liquid
stream and a second liquid stream; (b) said first liquid stream is
expanded to a first lower pressure and is thereafter supplied to a
first distillation column at an upper mid-column feed position; (c)
said second liquid stream is heated sufficiently to partially
vaporize it; (d) said partially vaporized second liquid stream is
separated thereby to provide a vapor stream and a third liquid
stream; (e) said vapor stream is expanded to said first lower
pressure and is supplied to said first distillation column at a
first lower mid-column feed position; (f) said third liquid stream
is expanded to said first lower pressure and thereafter supplied to
said first distillation column at a second lower mid-column feed
position; (g) a first overhead distillation stream is withdrawn
from an upper region of said first distillation column and
compressed to higher pressure; (h) said compressed first overhead
distillation stream is cooled sufficiently to at least partially
condense it and form thereby a condensed stream, with said cooling
supplying at least a portion of said heating of said second liquid
stream; (i) said condensed stream is divided into at least a
volatile liquid stream and a reflux liquid stream; (j) said reflux
liquid stream is further cooled, with said cooling supplying at
least a portion of said heating of said second liquid stream; (k)
said further cooled reflux liquid stream is divided into at least a
first reflux stream and a second reflux stream; (l) said first
reflux stream is supplied to said first distillation column at a
top column feed position; (m) said gas stream is divided into at
least a first gaseous stream and a second gaseous stream; (n) said
first gaseous stream is cooled to condense substantially all of it
and is thereafter expanded to a second lower pressure whereby it is
further cooled; (o) said expanded substantially condensed first
gaseous stream is thereafter supplied to a second distillation
column at an upper mid-column feed position; (p) said second
gaseous stream is expanded to said second lower pressure, is
cooled, and is thereafter supplied to said second distillation
column at a lower mid-column feed position; (q) said second reflux
stream is supplied to said second distillation column at a top
column feed position; (r) a second overhead distillation stream is
withdrawn from an upper region of said second distillation column;
(s) said volatile liquid stream is heated sufficiently to vaporize
it, with said heating supplying at least a portion of said cooling
of one or more of said first gaseous stream and said expanded
second gaseous stream; (t) said second overhead distillation stream
is heated, with said heating supplying at least a portion of said
cooling of one or more of said first gaseous stream and said
expanded second gaseous stream; (u) said vaporized volatile liquid
stream and said heated second overhead distillation stream are
combined to form said volatile residue gas fraction containing a
major portion of said methane; (v) a first bottom liquid from said
first distillation column and a second bottom liquid from said
second distillation column are combined to form said relatively
less volatile liquid fraction; and (w) the quantities and
temperatures of said first and second reflux streams and the
temperatures of said feeds to said first and second distillation
columns are effective to maintain the overhead temperatures of said
first and second distillation columns at temperatures whereby the
major portion of said heavier hydrocarbon components is recovered
in said relatively less volatile liquid fraction by fractionation
in said first and second distillation columns.
15. A process for the separation of liquefied natural gas
containing methane and heavier hydrocarbon components and a gas
stream containing methane and heavier hydrocarbon components into a
volatile residue gas fraction containing a major portion of said
methane and a relatively less volatile liquid fraction containing a
major portion of said heavier hydrocarbon components wherein (a)
said liquefied natural gas is divided into at least a first liquid
stream and a second liquid stream; (b) said first liquid stream is
expanded to a first lower pressure and is thereafter supplied to a
first distillation column at an upper mid-column feed position; (c)
said second liquid stream is heated sufficiently to partially
vaporize it; (d) said partially vaporized second liquid stream is
separated thereby to provide a first vapor stream and a third
liquid stream; (e) said first vapor stream is expanded to said
first lower pressure and thereafter supplied to said first
distillation column at a first lower mid-column feed position; (f)
said third liquid stream is expanded to said first lower pressure
and thereafter supplied to said first distillation column at a
second lower mid-column feed position; (g) a first overhead
distillation stream is withdrawn from an upper region of said first
distillation column and compressed to higher pressure; (h) said
compressed first overhead distillation stream is cooled
sufficiently to at least partially condense it and form thereby a
condensed stream, with said cooling supplying at least a portion of
said heating of said second liquid stream; (i) said condensed
stream is divided into at least a volatile liquid stream and a
reflux liquid stream; (j) said reflux liquid stream is further
cooled, with said cooling supplying at least a portion of said
heating of said second liquid stream; (k) said further cooled
reflux liquid stream is divided into at least a first reflux stream
and a second reflux stream; (l) said first reflux stream is
supplied to said first distillation column at a top column feed
position; (m) said gas stream is divided into at least a first
gaseous stream and a second gaseous stream; (n) said first gaseous
stream is cooled to condense substantially all of it and is
thereafter expanded to a second lower pressure whereby it is
further cooled; (o) said expanded substantially condensed first
gaseous stream is thereafter supplied to a second distillation
column at an upper mid-column feed position; (p) said second
gaseous stream is expanded to said second lower pressure and is
thereafter cooled sufficiently to partially condense it; (q) said
partially condensed expanded second gaseous stream is separated
thereby to provide a second vapor stream and a fourth liquid
stream; (r) said second vapor stream is further cooled and
thereafter supplied to said second distillation column at a first
lower mid-column feed position; (s) said fourth liquid stream is
supplied to said second distillation column at a second lower
mid-column feed position; (t) said second reflux stream is supplied
to said second distillation column at a top column feed position;
(u) a second overhead distillation stream is withdrawn from an
upper region of said second distillation column; (v) said volatile
liquid stream is heated sufficiently to vaporize it, with said
heating supplying at least a portion of said cooling of one or more
of said first gaseous stream, said expanded second gaseous stream,
and said second vapor stream; (w) said second overhead distillation
stream is heated, with said heating supplying at least a portion of
said cooling of one or more of said first gaseous stream, said
expanded second gaseous stream, and said second vapor stream; (x)
said vaporized volatile liquid stream and said heated second
overhead distillation stream are combined to form said volatile
residue gas fraction containing a major portion of said methane;
(y) a first bottom liquid from said first distillation column and a
second bottom liquid from said second distillation column are
combined to form said relatively less volatile liquid fraction; and
(z) the quantities and temperatures of said first and second reflux
streams and the temperatures of said feeds to said first and second
distillation columns are effective to maintain the overhead
temperatures of said first and second distillation columns at
temperatures whereby the major portion of said heavier hydrocarbon
components is recovered in said relatively less volatile liquid
fraction by fractionation in said first and second distillation
columns.
16. The process according to claim 12 or 14 wherein (a) said second
overhead distillation stream is compressed to higher pressure; (b)
said compressed second overhead distillation stream is heated, with
said heating supplying at least a portion of said cooling of one or
more of said first gaseous stream and said expanded second gaseous
stream; and (c) said vaporized volatile liquid stream and said
heated compressed second overhead distillation stream are combined
to form said volatile residue gas fraction.
17. The process according to claim 13 or 15 wherein (a) said second
overhead distillation stream is compressed to higher pressure; (b)
said compressed second overhead distillation stream is heated, with
said heating supplying at least a portion of said cooling of one or
more of said first gaseous stream, said expanded second gaseous
stream, and said second vapor stream; and (c) said vaporized
volatile liquid stream and said heated compressed second overhead
distillation stream are combined to form said volatile residue gas
fraction.
18. The process according to claim 12 or 14 wherein (a) said second
gaseous stream is cooled prior to said expansion; (b) said second
overhead distillation stream is compressed to higher pressure; (c)
said volatile liquid stream is heated sufficiently to vaporize it,
with said heating supplying at least a portion of said cooling of
one or more of said first gaseous stream, said second gaseous
stream, and said expanded second gaseous stream; (d) said
compressed second overhead distillation stream is heated, with said
heating supplying at least a portion of said cooling of one or more
of said first gaseous stream, said second gaseous stream, and said
expanded second gaseous stream; and (e) said vaporized volatile
liquid stream and said heated compressed second overhead
distillation stream are combined to form said volatile residue gas
fraction.
19. The process according to claim 13 wherein (a) said second
gaseous stream is cooled sufficiently to partially condense it; (b)
said partially condensed second gaseous stream is separated thereby
to provide said second vapor stream and said third liquid stream;
(c) said second vapor stream is expanded to said second lower
pressure, is cooled, and is thereafter supplied to said second
distillation column at said first lower mid-column feed position;
(d) said third liquid stream is expanded to said second lower
pressure and thereafter supplied to said second distillation column
at said second lower mid-column feed position; (e) said second
overhead distillation stream is compressed to higher pressure; (f)
said volatile liquid stream is heated sufficiently to vaporize it,
with said heating supplying at least a portion of said cooling of
one or more of said first gaseous stream, said second gaseous
stream, and said expanded second vapor stream; (g) said compressed
second overhead distillation stream is heated, with said heating
supplying at least a portion of said cooling of one or more of said
first gaseous stream, said second gaseous stream, and said expanded
second vapor stream; and (h) said vaporized volatile liquid stream
and said heated compressed second overhead distillation stream are
combined to form said volatile residue gas fraction.
20. The process according to claim 15 wherein (a) said second
gaseous stream is cooled sufficiently to partially condense it; (b)
said partially condensed second gaseous stream is separated thereby
to provide said second vapor stream and said fourth liquid stream;
(c) said second vapor stream is expanded to said second lower
pressure, is cooled, and is thereafter supplied to said second
distillation column at said first lower mid-column feed position;
(d) said fourth liquid stream is expanded to said second lower
pressure and thereafter supplied to said second distillation column
at said second lower mid-column feed position; (e) said second
overhead distillation stream is compressed to higher pressure; (f)
said volatile liquid stream is heated sufficiently to vaporize it,
with said heating supplying at least a portion of said cooling of
one or more of said first gaseous stream, said second gaseous
stream, and said expanded second vapor stream; (g) said compressed
second overhead distillation stream is heated, with said heating
supplying at least a portion of said cooling of one or more of said
first gaseous stream, said second gaseous stream, and said expanded
second vapor stream; and (h) said vaporized volatile liquid stream
and said heated compressed second overhead distillation stream are
combined to form said volatile residue gas fraction.
21. The process according to claim 19 wherein (a) said gas stream
is cooled sufficiently to partially condense it; (b) said partially
condensed gas stream is separated thereby to provide said second
vapor stream and said third liquid stream; (c) said second vapor
stream is divided into at least said first gaseous stream and said
second gaseous stream; (d) said second gaseous stream is expanded
to said second lower pressure, is cooled, and is thereafter
supplied to said second distillation column at said first lower
mid-column feed position; (e) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of one or more of said gas stream, said
first gaseous stream, and said expanded second gaseous stream; and
(f) said compressed second overhead distillation stream is heated,
with said heating supplying at least a portion of said cooling of
one or more of said gas stream, said first gaseous stream, and said
expanded second gaseous stream.
22. The process according to claim 20 wherein (a) said gas stream
is cooled sufficiently to partially condense it; (b) said partially
condensed gas stream is separated thereby to provide said second
vapor stream and said fourth liquid stream; (c) said second vapor
stream is divided into at least said first gaseous stream and said
second gaseous stream; (d) said second gaseous stream is expanded
to said second lower pressure, is cooled, and is thereafter
supplied to said second distillation column at said first lower
mid-column feed position; (e) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of one or more of said gas stream, said
first gaseous stream, and said expanded second gaseous stream; and
(f) said compressed second overhead distillation stream is heated,
with said heating supplying at least a portion of said cooling of
one or more of said gas stream, said first gaseous stream, and said
expanded second gaseous stream.
23. The process according to claim 1, 2, 3, 4, 8, 9, 10, or 11
wherein (a) said liquefied natural gas is heated and thereafter
divided into at least said first liquid stream and said second
liquid stream; and (b) said cooling of said compressed first
portion and said reflux stream supply at least a portion of said
heating of said liquefied natural gas.
24. The process according to claim 5 wherein (a) said liquefied
natural gas is heated and thereafter divided into at least said
first liquid stream and said second liquid stream; and (b) said
cooling of said compressed first portion and said reflux stream
supply at least a portion of said heating of said liquefied natural
gas.
25. The process according to claim 6 wherein (a) said liquefied
natural gas is heated and thereafter divided into at least said
first liquid stream and said second liquid stream; and (b) said
cooling of said compressed first portion and said reflux stream
supply at least a portion of said heating of said liquefied natural
gas.
26. The process according to claim 7 wherein (a) said liquefied
natural gas is heated and thereafter divided into at least said
first liquid stream and said second liquid stream; and (b) said
cooling of said compressed first portion and said reflux stream
supply at least a portion of said heating of said liquefied natural
gas.
27. The process according to claim 12, 13, 14, 15, 19, 20, 21, or
22 wherein (a) said liquefied natural gas is heated and thereafter
divided into at least said first liquid stream and said second
liquid stream; and (b) said cooling of said compressed first
overhead distillation stream and said reflux liquid stream supply
at least a portion of said heating of said liquefied natural
gas.
28. The process according to claim 16 wherein (a) said liquefied
natural gas is heated and thereafter divided into at least said
first liquid stream and said second liquid stream; and (b) said
cooling of said compressed first overhead distillation stream and
said reflux liquid stream supply at least a portion of said heating
of said liquefied natural gas.
29. The process according to claim 17 wherein (a) said liquefied
natural gas is heated and thereafter divided into at least said
first liquid stream and said second liquid stream; and (b) said
cooling of said compressed first overhead distillation stream and
said reflux liquid stream supply at least a portion of said heating
of said liquefied natural gas.
30. The process according to claim 18 wherein (a) said liquefied
natural gas is heated and thereafter divided into at least said
first liquid stream and said second liquid stream; and (b) said
cooling of said compressed first overhead distillation stream and
said reflux liquid stream supply at least a portion of said heating
of said liquefied natural gas.
31. The process according to claim 1, 2, 3, 4, 8, 9, 10, 11, 12,
13, 14, 15, 19, 20, 21, or 22 wherein said volatile residue gas
fraction contains a major portion of said methane and C.sub.2
components.
32. The process according to claim 5 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
33. The process according to claim 6 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
34. The process according to claim 7 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
35. The process according to claim 16 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components
36. The process according to claim 17 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
37. The process according to claim 18 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
38. The process according to claim 23 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
39. The process according to claim 24 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
40. The process according to claim 25 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
41. The process according to claim 26 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
42. The process according to claim 27 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
43. The process according to claim 28 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
44. The process according to claim 29 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
45. The process according to claim 30 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
Description
[0001] This invention relates to a process for the separation of
ethane and heavier hydrocarbons or propane and heavier hydrocarbons
from liquefied natural gas (hereinafter referred to as LNG)
combined with the separation of a gas containing hydrocarbons to
provide a volatile methane-rich gas stream and a less volatile
natural gas liquids (NGL) or liquefied petroleum gas (LPG) stream.
The applicants claim the benefits under Title 35, United States
Code, Section 119(e) of prior U.S. Provisional Application No.
61/053,814 which was filed on May 16, 2008.
BACKGROUND OF THE INVENTION
[0002] As an alternative to transportation in pipelines, natural
gas at remote locations is sometimes liquefied and transported in
special LNG tankers to appropriate LNG receiving and storage
terminals. The LNG can then be re-vaporized and used as a gaseous
fuel in the same fashion as natural gas. Although LNG usually has a
major proportion of methane, i.e., methane comprises at least 50
mole percent of the LNG, it also contains relatively lesser amounts
of heavier hydrocarbons such as ethane, propane, butanes, and the
like, as well as nitrogen. It is often necessary to separate some
or all of the heavier hydrocarbons from the methane in the LNG so
that the gaseous fuel resulting from vaporizing the LNG conforms to
pipeline specifications for heating value. In addition, it is often
also desirable to separate the heavier hydrocarbons from the
methane and ethane because these hydrocarbons have a higher value
as liquid products (for use as petrochemical feedstocks, as an
example) than their value as fuel.
[0003] Although there are many processes which may be used to
separate ethane and/or propane and heavier hydrocarbons from LNG,
these processes often must compromise between high recovery, low
utility costs, and process simplicity (and hence low capital
investment). U.S. Pat. Nos. 2,952,984; 3,837,172; 5,114,451; and
7,155,931 describe relevant LNG processes capable of ethane or
propane recovery while producing the lean LNG as a vapor stream
that is thereafter compressed to delivery pressure to enter a gas
distribution network. However, lower utility costs may be possible
if the lean LNG is instead produced as a liquid stream that can be
pumped (rather than compressed) to the delivery pressure of the gas
distribution network, with the lean LNG subsequently vaporized
using a low level source of external heat or other means. U.S. Pat.
Nos. 6,604,380; 6,907,752; 6,941,771; 7,069,743; and 7,216,507 and
co-pending application Ser. Nos. 11/749,268 and 12/060,362 describe
such processes.
[0004] Economics and logistics often dictate that LNG receiving
terminals be located close to the natural gas transmission lines
that will transport the re-vaporized LNG to consumers. In many
cases, these areas also have plants for processing natural gas
produced in the region to recover the heavier hydrocarbons
contained in the natural gas. Available processes for separating
these heavier hydrocarbons include those based upon cooling and
refrigeration of gas, oil absorption, and refrigerated oil
absorption. Additionally, cryogenic processes have become popular
because of the availability of economical equipment that produces
power while simultaneously expanding and extracting heat from the
gas being processed. Depending upon the pressure of the gas source,
the richness (ethane, ethylene, and heavier hydrocarbons content)
of the gas, and the desired end products, each of these processes
or a combination thereof may be employed.
[0005] The cryogenic expansion process is now generally preferred
for natural gas liquids recovery because it provides maximum
simplicity with ease of startup, operating flexibility, good
efficiency, safety, and good reliability. U.S. Pat. Nos. 3,292,380;
4,061,481; 4,140,504; 4,157,904; 4,171,964; 4,185,978; 4,251,249;
4,278,457; 4,519,824; 4,617,039; 4,687,499; 4,689,063; 4,690,702;
4,854,955; 4,869,740; 4,889,545; 5,275,005; 5,555,748; 5,566,554;
5,568,737; 5,771,712; 5,799,507; 5,881,569; 5,890,378; 5,983,664;
6,182,469; 6,578,379; 6,712,880; 6,915,662; 7,191,617; 7,219,513;
reissue U.S. Pat. No. 33,408; and co-pending application Ser. Nos.
11/430,412; 11/839,693; 11/971,491; and 12/206,230 describe
relevant processes (although the description of the present
invention is based on different processing conditions than those
described in the cited U.S. patents).
[0006] The present invention is generally concerned with the
integrated recovery of ethylene, ethane, propylene, propane, and
heavier hydrocarbons from such LNG and gas streams. It uses a novel
process arrangement to integrate the heating of the LNG stream and
the cooling of the gas stream to eliminate the need for a separate
vaporizer and the need for external refrigeration, allowing high
C.sub.2 component recovery while keeping the processing equipment
simple and the capital investment low. Further, the present
invention offers a reduction in the utilities (power and heat)
required to process the LNG and gas streams, resulting in lower
operating costs than other processes, and also offering significant
reduction in capital investment.
[0007] Heretofore, assignee's U.S. Pat. No. 7,216,507 has been used
to recover C.sub.2 components and heavier hydrocarbon components in
plants processing LNG, while assignee's U.S. Pat. No. 5,568,737 has
been used to recover C.sub.2 components and heavier hydrocarbon
components in plants processing natural gas. Surprisingly,
applicants have found that by integrating certain features of the
assignee's U.S. Pat. No. 7,216,507 invention with certain features
of the assignee's U.S. Pat. No. 5,568,737, extremely high C.sub.2
component recovery levels can be accomplished using less energy
than that required by individual plants to process the LNG and
natural gas separately.
[0008] A typical analysis of an LNG stream to be processed in
accordance with this invention would be, in approximate mole
percent, 92.2% methane, 6.0% ethane and other C.sub.2 components,
1.1% propane and other C.sub.3 components, and traces of butanes
plus, with the balance made up of nitrogen. A typical analysis of a
gas stream to be processed in accordance with this invention would
be, in approximate mole percent, 80.1% methane, 9.5% ethane and
other C.sub.2 components, 5.6% propane and other C.sub.3
components, 1.3% iso-butane, 1.1% normal butane, 0.8% pentanes
plus, with the balance made up of nitrogen and carbon dioxide.
Sulfur containing gases are also sometimes present.
[0009] For a better understanding of the present invention,
reference is made to the following examples and drawings. Referring
to the drawings:
[0010] FIG. 1 is a flow diagram of a base case natural gas
processing plant using LNG to provide its refrigeration;
[0011] FIG. 2 is a flow diagram of base case LNG and natural gas
processing plants in accordance with U.S. Pat. Nos. 7,216,507 and
5,568,737, respectively;
[0012] FIG. 3 is a flow diagram of an LNG and natural gas
processing plant in accordance with the present invention; and
[0013] FIGS. 4 through 8 are flow diagrams illustrating alternative
means of application of the present invention to LNG and natural
gas streams.
[0014] FIGS. 1 and 2 are provided to quantify the advantages of the
present invention.
[0015] In the following explanation of the above figures, tables
are provided summarizing flow rates calculated for representative
process conditions. In the tables appearing herein, the values for
flow rates (in moles per hour) have been rounded to the nearest
whole number for convenience. The total stream rates shown in the
tables include all non-hydrocarbon components and hence are
generally larger than the sum of the stream flow rates for the
hydrocarbon components. Temperatures indicated are approximate
values rounded to the nearest degree. It should also be noted that
the process design calculations performed for the purpose of
comparing the processes depicted in the figures are based on the
assumption of no heat leak from (or to) the surroundings to (or
from) the process. The quality of commercially available insulating
materials makes this a very reasonable assumption and one that is
typically made by those skilled in the art.
[0016] For convenience, process parameters are reported in both the
traditional British units and in the units of the Systeme
International d'Unites (SI). The molar flow rates given in the
tables may be interpreted as either pound moles per hour or
kilogram moles per hour. The energy consumptions reported as
horsepower (HP) and/or thousand British Thermal Units per hour
(MBTU/Hr) correspond to the stated molar flow rates in pound moles
per hour. The energy consumptions reported as kilowatts (kW)
correspond to the stated molar flow rates in kilogram moles per
hour.
[0017] FIG. 1 is a flow diagram showing the design of a processing
plant to recover C.sub.2+ components from natural gas using an LNG
stream to provide refrigeration. In the simulation of the FIG. 1
process, inlet gas enters the plant at 126.degree. F. [52.degree.
C.] and 600 psia [4,137 kPa(a)] as stream 31. If the inlet gas
contains a concentration of sulfur compounds which would prevent
the product streams from meeting specifications, the sulfur
compounds are removed by appropriate pretreatment of the feed gas
(not illustrated). In addition, the feed stream is usually
dehydrated to prevent hydrate (ice) formation under cryogenic
conditions. Solid desiccant has typically been used for this
purpose.
[0018] The inlet gas stream 31 is cooled in heat exchanger 12 by
heat exchange with a portion (stream 72a) of partially warmed LNG
at -174.degree. F. [-114.degree. C.] and cool distillation stream
38a at -107.degree. F. [-77.degree. C.]. The cooled stream 31a
enters separator 13 at -79.degree. F. [-62.degree. C.] and 584 psia
[4,027 kPa(a)] where the vapor (stream 34) is separated from the
condensed liquid (stream 35). Liquid stream 35 is flash expanded
through an appropriate expansion device, such as expansion valve
17, to the operating pressure (approximately 430 psia [2,965
kPa(a)]) of fractionation tower 20. The expanded stream 35a leaving
expansion valve 17 reaches a temperature of -93.degree. F.
[-70.degree. C.] and is supplied to fractionation tower 20 at a
first mid-column feed point.
[0019] The vapor from separator 13 (stream 34) enters a work
expansion machine 10 in which mechanical energy is extracted from
this portion of the high pressure feed. The machine 10 expands the
vapor substantially isentropically to slightly above the tower
operating pressure, with the work expansion cooling the expanded
stream 34a to a temperature of approximately -101.degree. F.
[-74.degree. C.]. The typical commercially available expanders are
capable of recovering on the order of 80-88% of the work
theoretically available in an ideal isentropic expansion. The work
recovered is often used to drive a centrifugal compressor (such as
item 11) that can be used to re-compress the heated distillation
stream (stream 38b), for example. The expanded stream 34a is
further cooled to -124.degree. F. [-87.degree. C.] in heat
exchanger 14 by heat exchange with cold distillation stream 38 at
-143.degree. F. [-97.degree. C.], whereupon the partially condensed
expanded stream 34b is thereafter supplied to fractionation tower
20 at a second mid-column feed point.
[0020] The demethanizer in tower 20 is a conventional distillation
column containing a plurality of vertically spaced trays, one or
more packed beds, or some combination of trays and packing to
provide the necessary contact between the liquids falling downward
and the vapors rising upward. The column also includes reboilers
(such as reboiler 19) which heat and vaporize a portion of the
liquids flowing down the column to provide the stripping vapors
which flow up the column to strip the liquid product, stream 41, of
methane and lighter components. Liquid product stream 41 exits the
bottom of the tower at 99.degree. F. [37.degree. C.], based on a
typical specification of a methane to ethane ratio of 0.020:1 on a
molar basis in the bottom product.
[0021] Overhead distillation stream 43 is withdrawn from the upper
section of fractionation tower 20 at -143.degree. F. [-97.degree.
C.] and is divided into two portions, streams 44 and 47. The first
portion, stream 44, flows to reflux condenser 22 where it is cooled
to -237.degree. F. [-149.degree. C.] and totally condensed by heat
exchange with a portion (stream 72) of the cold LNG (stream 71a).
Condensed stream 44a enters reflux separator 23 wherein the
condensed liquid (stream 46) is separated from any uncondensed
vapor (stream 45). The liquid stream 46 from reflux separator 23 is
pumped by reflux pump 24 to a pressure slightly above the operating
pressure of demethanizer 20 and stream 46a is then supplied as cold
top column feed (reflux) to demethanizer 20. This cold liquid
reflux absorbs and condenses the C.sub.2 components and heavier
hydrocarbon components from the vapors rising in the upper section
of demethanizer 20.
[0022] The second portion (stream 47) of overhead vapor stream 43
combines with any uncondensed vapor (stream 45) from reflux
separator 23 to form cold distillation stream 38 at -143.degree. F.
[-97.degree. C.]. Distillation stream 38 passes countercurrently to
expanded stream 34a in heat exchanger 14 where it is heated to
-107.degree. F. [-77.degree. C.] (stream 38a), and countercurrently
to inlet gas in heat exchanger 12 where it is heated to 47.degree.
F. [8.degree. C.] (stream 38b). The distillation stream is then
re-compressed in two stages. The first stage is compressor 11
driven by expansion machine 10. The second stage is compressor 21
driven by a supplemental power source which compresses stream 38c
to sales line pressure (stream 38d). After cooling to 126.degree.
F. [52.degree. C.] in discharge cooler 22, stream 38e combines with
warm LNG stream 71b to form the residue gas product (stream 42).
Residue gas stream 42 flows to the sales gas pipeline at 1262 psia
[8,701 kPa(a)], sufficient to meet line requirements.
[0023] The LNG (stream 71) from LNG tank 50 enters pump 51 at
-251.degree. F. [-157.degree. C.]. Pump 51 elevates the pressure of
the LNG sufficiently so that it can flow through heat exchangers
and thence to the sales gas pipeline. Stream 71a exits the pump 51
at -242.degree. F. [-152.degree. C.] and 1364 psia [9,401 kPa(a)]
and is divided into two portions, streams 72 and 73. The first
portion, stream 72, is heated as described previously to
-174.degree. F. [-114.degree. C.] in reflux condenser 22 as it
provides cooling to the portion (stream 44) of overhead vapor
stream 43 from fractionation tower 20, and to 43.degree. F.
[6.degree. C.] in heat exchanger 12 as it provides cooling to the
inlet gas. The second portion, stream 73, is heated to 35.degree.
F. [2.degree. C.] in heat exchanger 53 using low level utility
heat. The heated streams 72b and 73a recombine to form warm LNG
stream 71b at 40.degree. F. [4.degree. C.], which thereafter
combines with distillation stream 38e to form residue gas stream 42
as described previously.
[0024] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 1 is set forth in the following
table:
TABLE-US-00001 TABLE I (FIG. 1) Stream Flow Summary - Lb. Moles/Hr
[kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31
42,545 5,048 2,972 1,658 53,145 34 33,481 1,606 279 39 36,221 35
9,064 3,442 2,693 1,619 16,924 43 50,499 25 0 0 51,534 44 8,055 4 0
0 8,221 45 0 0 0 0 0 46 8,055 4 0 0 8,221 47 42,444 21 0 0 43,313
38 42,444 21 0 0 43,313 71 40,293 2,642 491 3 43,689 72 27,601
1,810 336 2 29,927 73 12,692 832 155 1 13,762 42 82,737 2,663 491 3
87,002 41 101 5,027 2,972 1,658 9,832 Recoveries* Ethane 65.37%
Propane 85.83% Butanes+ 99.83% Power LNG Feed Pump 3,561 HP [ 5,854
kW] Reflux Pump 23 HP [ 38 kW] Residue Gas Compressor 24,612 HP [
40,462 kW] Totals 28,196 HP [ 46,354 kW] Low Level Utility Heat LNG
Heater 68,990 MBTU/Hr [ 44,564 kW] High Level Utility Heat
Demethanizer Reboiler 80,020 MBTU/Hr [ 51,689 kW] Specific Power
HP-Hr/Lb. Mole 2.868 [kW-Hr/kg mole] [ 4.715 ] *(Based on
un-rounded flow rates)
[0025] The recoveries reported in Table I are computed relative to
the total quantities of ethane, propane, and butanes+ contained in
the gas stream being processed in the plant and in the LNG stream.
Although the recoveries are quite high relative to the heavier
hydrocarbons contained in the gas being processed (99.58%, 100.00%,
and 100.00%, respectively, for ethane, propane, and butanes+), none
of the heavier hydrocarbons contained in the LNG stream are
captured in the FIG. 1 process. In fact, depending on the
composition of LNG stream 71, the residue gas stream 42 produced by
the FIG. 1 process may not meet all pipeline specifications. The
specific power reported in Table I is the power consumed per unit
of liquid product recovered, and is an indicator of the overall
process efficiency.
[0026] FIG. 2 is a flow diagram showing processes to recover
C.sub.2+ components from LNG and natural gas in accordance with
U.S. Pat. Nos. 7,216,507 and 5,568,737, respectively, with the
processed LNG stream used to provide refrigeration for the natural
gas plant. The processes of FIG. 2 have been applied to the same
LNG stream and inlet gas stream compositions and conditions as
described previously for FIG. 1.
[0027] In the simulation of the FIG. 2 process, the LNG to be
processed (stream 71) from LNG tank 50 enters pump 51 at
-251.degree. F. [-157.degree. C.]. Pump 51 elevates the pressure of
the LNG sufficiently so that it can flow through heat exchangers
and thence to expansion machine 55. Stream 71a exits the pump at
-242.degree. F. [-152.degree. C.] and 1364 psia [9,401 kPa(a)] and
is split into two portions, streams 75 and 76. The first portion,
stream 75, is expanded to the operating pressure (approximately 415
psia [2,859 kPa(a)]) of fractionation column 62 by expansion valve
58. The expanded stream 75a leaves expansion valve 58 at
-238.degree. F. [-150.degree. C.] and is thereafter supplied to
tower 62 at an upper mid-column feed point.
[0028] The second portion, stream 76, is heated to -79.degree. F.
[-62.degree. C.] in heat exchanger 52 by cooling compressed
overhead distillation stream 79a at -70.degree. F. [-57.degree. C.]
and reflux stream 82 at -128.degree. F. [-89.degree. C.]. The
partially heated stream 76a is further heated and vaporized in heat
exchanger 53 using low level utility heat. The heated stream 76b at
-5.degree. F. [-20.degree. C.] and 1334 psia [9,195 kPa(a)] enters
work expansion machine 55 in which mechanical energy is extracted
from this portion of the high pressure feed. The machine 55 expands
the vapor substantially isentropically to the tower operating
pressure, with the work expansion cooling the expanded stream 76c
to a temperature of approximately -107.degree. F. [-77.degree. C.]
before it is supplied as feed to fractionation column 62 at a lower
mid-column feed point.
[0029] The demethanizer in fractionation column 62 is a
conventional distillation column containing a plurality of
vertically spaced trays, one or more packed beds, or some
combination of trays and packing consisting of two sections. The
upper absorbing (rectification) section contains the trays and/or
packing to provide the necessary contact between the vapors rising
upward and cold liquid falling downward to condense and absorb the
ethane and heavier components; the lower stripping (demethanizing)
section contains the trays and/or packing to provide the necessary
contact between the liquids falling downward and the vapors rising
upward. The demethanizing section also includes one or more
reboilers (such as side reboiler 60 using low level utility heat,
and reboiler 61 using high level utility heat) which heat and
vaporize a portion of the liquids flowing down the column to
provide the stripping vapors which flow up the column. The column
liquid stream 80 exits the bottom of the tower at 54.degree. F.
[12.degree. C.], based on a typical specification of a methane to
ethane ratio of 0.020:1 on a molar basis in the bottom product.
[0030] Overhead distillation stream 79 is withdrawn from the upper
section of fractionation tower 62 at -144.degree. F. [-98.degree.
C.] and flows to compressor 56 driven by expansion machine 55,
where it is compressed to 807 psia [5,567 kPa(a)] (stream 79a). At
this pressure, the stream is totally condensed as it is cooled to
-128.degree. F. [-89.degree. C.] in heat exchanger 52 as described
previously. The condensed liquid (stream 79b) is then divided into
two portions, streams 83 and 82. The first portion (stream 83) is
the methane-rich lean LNG stream, which is pumped by pump 63 to
1270 psia [8,756 kPa(a)] for subsequent vaporization in heat
exchanger 12, heating stream 83a to 40.degree. F. [4.degree. C.] as
described in paragraph [0032] below to produce warm lean LNG stream
83b.
[0031] The remaining portion of condensed liquid stream 79b, reflux
stream 82, flows to heat exchanger 52 where it is subcooled to
-237.degree. F. [-149.degree. C.] by heat exchange with a portion
of the cold LNG (stream 76) as described previously. The subcooled
stream 82a is then expanded to the operating pressure of
demethanizer 62 by expansion valve 57. The expanded stream 82b at
-236.degree. F. [-149.degree. C.] is then supplied as cold top
column feed (reflux) to demethanizer 62. This cold liquid reflux
absorbs and condenses the C.sub.2 components and heavier
hydrocarbon components from the vapors rising in the upper
rectification section of demethanizer 62.
[0032] In the simulation of the FIG. 2 process, inlet gas enters
the plant at 126.degree. F. [52.degree. C.] and 600 psia [4,137
kPa(a)] as stream 31. The feed stream 31 is cooled in heat
exchanger 12 by heat exchange with cold lean LNG (stream 83a) at
-116.degree. F. [-82.degree. C.], cool distillation stream 38a at
-96.degree. F. [-71.degree. C.], and demethanizer liquids (stream
39) at -3.degree. F. [-20.degree. C.]. The cooled stream 31a enters
separator 13 at -67.degree. F. [-55.degree. C.] and 584 psia [4,027
kPa(a)] where the vapor (stream 33) is separated from the condensed
liquid (stream 35). Liquid stream 35 is flash expanded through an
appropriate expansion device, such as expansion valve 17, to the
operating pressure (approximately 375 psia [2,583 kPa(a)]) of
fractionation tower 20. The expanded stream 35a leaving expansion
valve 17 reaches a temperature of -86.degree. F. [-65.degree. C.]
and is supplied to fractionation tower 20 at a first lower
mid-column feed point.
[0033] Vapor stream 33 from separator 13 is divided into two
streams, 32 and 34. Stream 32, containing about 22% of the total
vapor, passes through heat exchanger 14 in heat exchange relation
with cold distillation stream 38 at -150.degree. F. [-101.degree.
C.] where it is cooled to substantial condensation. The resulting
substantially condensed stream 32a at -144.degree. F. [-98.degree.
C.] is then flash expanded through an appropriate expansion device,
such as expansion valve 16, to the operating pressure of
fractionation tower 20, cooling stream 32b to -148.degree. F.
[-100.degree. C.] before it is supplied to fractionation tower 20
at an upper mid-column feed point.
[0034] The remaining 78% of the vapor from separator 13 (stream 34)
enters a work expansion machine 10 in which mechanical energy is
extracted from this portion of the high pressure feed. The machine
10 expands the vapor substantially isentropically to the tower
operating pressure, with the work expansion cooling the expanded
stream 34a to a temperature of approximately -100.degree. F.
[-73.degree. C.]. The partially condensed expanded stream 34a is
thereafter supplied as feed to fractionation tower 20 at a second
lower mid-column feed point.
[0035] The demethanizer in fractionation column 20 is a
conventional distillation column containing a plurality of
vertically spaced trays, one or more packed beds, or some
combination of trays and packing consisting of two sections. The
upper absorbing (rectification) section contains the trays and/or
packing to provide the necessary contact between the vapors rising
upward and cold liquid falling downward to condense and absorb the
ethane and heavier components; the lower stripping (demethanizing)
section contains the trays and/or packing to provide the necessary
contact between the liquids falling downward and the vapors rising
upward. The demethanizing section also includes one or more
reboilers (such as the side reboiler in heat exchanger 12 described
previously, and reboiler 19 using high level utility heat) which
heat and vaporize a portion of the liquids flowing down the column
to provide the stripping vapors which flow up the column. The
column liquid stream 40 exits the bottom of the tower at 85.degree.
F. [30.degree. C.], based on a typical specification of a methane
to ethane ratio of 0.020:1 on a molar basis in the bottom product,
and combines with stream 80 to form the liquid product (stream
41).
[0036] Overhead distillation stream 38 is withdrawn from the upper
section of fractionation tower 20 at -150.degree. F. [-101.degree.
C.]. It passes countercurrently to vapor stream 32 and recycle
stream 36a in heat exchanger 14 where it is heated to -96.degree.
F. [-71.degree. C.] (stream 38a), and countercurrently to inlet gas
stream 31 and recycle stream 36 in heat exchanger 12 where it is
heated to 6.degree. F. [-15.degree. C.] (stream 38b). The
distillation stream is then re-compressed in two stages. The first
stage is compressor 11 driven by expansion machine 10. The second
stage is compressor 21 driven by a supplemental power source which
compresses stream 38c to sales line pressure (stream 38d). After
cooling to 126.degree. F. [52.degree. C.] in discharge cooler 22,
stream 38e is divided into two portions, stream 37 and recycle
stream 36. Stream 37 combines with warm lean LNG stream 83b to form
the residue gas product (stream 42). Residue gas stream 42 flows to
the sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient to
meet line requirements.
[0037] Recycle stream 36 flows to heat exchanger 12 and is cooled
to -102.degree. F. [-75.degree. C.] by heat exchange with cool lean
LNG (stream 83a), cool distillation stream 38a, and demethanizer
liquids (stream 39) as described previously. Stream 36a is further
cooled to -144.degree. F. [-98.degree. C.] by heat exchange with
cold distillation stream 38 in heat exchanger 14 as described
previously. The substantially condensed stream 36b is then expanded
through an appropriate expansion device, such as expansion valve
15, to the demethanizer operating pressure, resulting in cooling of
the total stream to -152.degree. F. [-102.degree. C.]. The expanded
stream 36c is then supplied to fractionation tower 20 as the top
column feed. The vapor portion of stream 36c combines with the
vapors rising from the top fractionation stage of the column to
form distillation stream 38, which is withdrawn from an upper
region of the tower as described above.
[0038] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 2 is set forth in the following
table:
TABLE-US-00002 TABLE II (FIG. 2) Stream Flow Summary - Lb. Moles/Hr
[kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31
42,545 5,048 2,972 1,658 53,145 33 36,197 2,152 429 64 39,690 35
6,348 2,896 2,543 1,594 13,455 32 8,027 477 95 14 8,801 34 28,170
1,675 334 50 30,889 38 52,982 30 0 0 54,112 36 10,537 6 0 0 10,762
37 42,445 24 0 0 43,350 40 100 5,024 2,972 1,658 9,795 71 40,293
2,642 491 3 43,689 75 4,835 317 59 0 5,243 76 35,458 2,325 432 3
38,446 79 45,588 16 0 0 45,898 82 5,348 2 0 0 5,385 83 40,240 14 0
0 40,513 80 53 2,628 491 3 3,176 42 82,685 38 0 0 83,863 41 153
7,652 3,463 1,661 12,971 Recoveries* Ethane 99.51% Propane 100.00%
Butanes+ 100.00% Power LNG Feed Pump 3,561 HP [ 5,854 kW] LNG
Product Pump 1,746 HP [ 2,870 kW] Residue Gas Compressor 31,674 HP
[ 52,072 kW] Totals 36,981 HP [ 60,796 kW] Low Level Utility Heat
Liquid Feed Heater 66,200 MBTU/Hr [ 42,762 kW] Demethanizer
Reboiler 60 23,350 MBTU/Hr [ 15,083 kW] Totals 89,550 MBTU/Hr [
57,845 kW] High Level Utility Heat Demethanizer Reboiler 19 20,080
MBTU/Hr [ 12,971 kW] Demethanizer Reboiler 61 3,400 MBTU/Hr [ 2,196
kW] Totals 23,480 MBTU/Hr [ 15,167 kW] Specific Power HP-Hr/Lb.
Mole 2.851 [kW-Hr/kg mole] [ 4.687 ] *(Based on un-rounded flow
rates)
[0039] Comparison of the recovery levels displayed in Tables I and
II shows that the liquids recovery of the FIG. 2 processes is much
higher than that of the FIG. 1 process due to the recovery of the
heavier hydrocarbon liquids contained in the LNG stream in
fractionation tower 62. The ethane recovery improves from 65.37% to
99.51%, the propane recovery improves from 85.83% to 100.00%, and
the butanes+ recovery improves from 99.83% to 100.00%. In addition,
the process efficiency of the FIG. 2 processes is improved by about
1% in terms of the specific power relative to the FIG. 1
process.
DESCRIPTION OF THE INVENTION
EXAMPLE 1
[0040] FIG. 3 illustrates a flow diagram of a process in accordance
with the present invention. The LNG stream and inlet gas stream
compositions and conditions considered in the process presented in
FIG. 3 are the same as those in the FIG. 1 and FIG. 2 processes.
Accordingly, the FIG. 3 process can be compared with the FIG. 1 and
FIG. 2 processes to illustrate the advantages of the present
invention.
[0041] In the simulation of the FIG. 3 process, the LNG to be
processed (stream 71) from LNG tank 50 enters pump 51 at
-251.degree. F. [-157.degree. C.]. Pump 51 elevates the pressure of
the LNG sufficiently so that it can flow through heat exchangers
and thence to separator 54. Stream 71a exits the pump at
-242.degree. F. [-152.degree. C.] and 1364 psia [9,401 kPa(a)] and
is split into two portions, streams 72 and 73. The first portion,
stream 72, becomes stream 75 and is expanded to the operating
pressure (approximately 415 psia [2,859 kPa(a)]) of fractionation
column 62 by expansion valve 58. The expanded stream 75a leaves
expansion valve 58 at -238.degree. F. [-150C.] and is thereafter
supplied to tower 62 at an upper mid-column feed point.
[0042] The second portion, stream 73, is heated prior to entering
separator 54 so that all or a portion of it is vaporized. In the
example shown in FIG. 3, stream 73 is first heated to -77.degree.
F. [-61 .degree. C.] in heat exchanger 52 by cooling compressed
overhead distillation stream 79a at -70.degree. F. [-57.degree. C.]
and reflux stream 81 at -16.degree. F. [-82.degree. C.]. The
partially heated stream 73a becomes stream 76 and is further heated
in heat exchanger 53 using low level utility heat. (High level
utility heat, such as the heating medium used in tower reboiler 61,
is normally more expensive than low level utility heat, so lower
operating cost is usually achieved when use of low level heat, such
as sea water, is maximized and the use of high level utility heat
is minimized.) Note that in all cases exchangers 52 and 53 are
representative of either a multitude of individual heat exchangers
or a single multi-pass heat exchanger, or any combination thereof.
(The decision as to whether to use more than one heat exchanger for
the indicated heating services will depend on a number of factors
including, but not limited to, inlet LNG flow rate, heat exchanger
size, stream temperatures, etc.)
[0043] The heated stream 76a enters separator 54 at -5.degree. F.
[-20.degree. C.] and 1334 psia [9,195 kPa(a)] where the vapor
(stream 77) is separated from any remaining liquid (stream 78).
Vapor stream 77 enters a work expansion machine 55 in which
mechanical energy is extracted from this portion of the high
pressure feed. The machine 55 expands the vapor substantially
isentropically to the tower operating pressure, with the work
expansion cooling the expanded stream 77a to a temperature of
approximately -107.degree. F. [-77.degree. C.]. The work recovered
is often used to drive a centrifugal compressor (such as item 56)
that can be used to re-compress the column overhead vapor (stream
79), for example. The partially condensed expanded stream 77a is
thereafter supplied as feed to fractionation column 62 at a lower
mid-column feed point. The separator liquid (stream 78), if any, is
expanded to the operating pressure of fractionation column 62 by
expansion valve 59 before expanded stream 78a is supplied to
fractionation tower 62 at a second lower mid-column feed point.
[0044] The demethanizer in fractionation column 62 is a
conventional distillation column containing a plurality of
vertically spaced trays, one or more packed beds, or some
combination of trays and packing. The fractionation tower 62 may
consist of two sections. The upper absorbing (rectification)
section contains the trays and/or packing to provide the necessary
contact between the vapors rising upward and cold liquid falling
downward to condense and absorb the ethane and heavier components;
the lower stripping (demethanizing) section contains the trays
and/or packing to provide the necessary contact between the liquids
falling downward and the vapors rising upward. The demethanizing
section also includes one or more reboilers (such as side reboiler
60 using low level utility heat, and reboiler 61 using high level
utility heat) which heat and vaporize a portion of the liquids
flowing down the column to provide the stripping vapors which flow
up the column. The column liquid stream 80 exits the bottom of the
tower at 54.degree. F. [12.degree. C.], based on a typical
specification of a methane to ethane ratio of 0.020:1 on a molar
basis in the bottom product.
[0045] Overhead distillation stream 79 is withdrawn from the upper
section of fractionation tower 62 at -144.degree. F. [-98.degree.
C.] and flows to compressor 56 driven by expansion machine 55,
where it is compressed to 805 psia [5,554 kPa(a)] (stream 79a). At
this pressure, the stream is totally condensed as it is cooled to
-116.degree. F. [-82.degree. C.] in heat exchanger 52 as described
previously. The condensed liquid (stream 79b) is then divided into
two portions, streams 83 and 81. The first portion (stream 83) is
the methane-rich lean LNG stream, which is pumped by pump 63 to
1275 psia [8,791 kPa(a)] for subsequent vaporization in heat
exchangers 14 and 12, heating stream 83a to -94.degree. F.
[-70.degree. C.] and 40.degree. F. [4.degree. C.], respectively, as
described in paragraphs [0047] and [0049] below to produce warm
lean LNG stream 83c.
[0046] The remaining portion of condensed liquid stream 79b, stream
81, flows to heat exchanger 52 where it is subcooled to
-237.degree. F. [-149.degree. C.] by heat exchange with a portion
of the cold LNG (stream 73) as described previously. The subcooled
stream 81a is then divided into two portions, streams 82 and 36.
The first portion, reflux stream 82, is expanded to the operating
pressure of demethanizer 62 by expansion valve 57. The expanded
stream 82a at -236.degree. F. [-149.degree. C.] is then supplied as
cold top column feed (reflux) to demethanizer 62. This cold liquid
reflux absorbs and condenses the C.sub.2 components and heavier
hydrocarbon components from the vapors rising in the upper
rectification section of demethanizer 62. The disposition of the
second portion, reflux stream 36 for demethanizer 20, is described
in paragraph [0050] below.
[0047] In the simulation of the FIG. 3 process, inlet gas enters
the plant at 126.degree. F. [52.degree. C.] and 600 psia [4,137
kPa(a)] as stream 31. The feed stream 31 is divided into two
portions, streams 32 and 33. The first portion, stream 32, is
cooled in heat exchanger 12 by heat exchange with cool lean LNG
(stream 83b) at -94.degree. F. [-70.degree. C.], cool distillation
stream 38a at -94.degree. F. [-70.degree. C.], and demethanizer
liquids (stream 39) at -78.degree. F. [-61.degree. C.]. The
partially cooled stream 32a is further cooled from -89.degree. F.
[-67.degree. C.] to -120.degree. F. [-85.degree. C.] in heat
exchanger 14 by heat exchange with cold lean LNG (stream 83a) at
-97.degree. F. [-72.degree. C.] and cold distillation stream 38 at
-144.degree. F. [-98.degree. C.]. Note that in all cases exchangers
12 and 14 are representative of either a multitude of individual
heat exchangers or a single multi-pass heat exchanger, or any
combination thereof. (The decision as to whether to use more than
one heat exchanger for the indicated heating services will depend
on a number of factors including, but not limited to, inlet gas
flow rate, heat exchanger size, stream temperatures, etc.) The
substantially condensed stream 32b is then flash expanded through
an appropriate expansion device, such as expansion valve 16, to the
operating pressure (approximately 415 psia [2,861 kPa(a)]) of
fractionation tower 20, cooling stream 32c to -132.degree. F.
[-91.degree. C.] before it is supplied to fractionation tower 20 at
an upper mid-column feed point.
[0048] The second portion of feed stream 31, stream 33, enters a
work expansion machine 10 in which mechanical energy is extracted
from this portion of the high pressure feed. The machine 10 expands
the vapor substantially isentropically to a pressure slightly above
the operating pressure of fractionation tower 20, with the work
expansion cooling the expanded stream 33a to a temperature of
approximately 92.degree. F. [33.degree. C.]. The work recovered is
often used to drive a centrifugal compressor (such as item 11) that
can be used to re-compress the heated distillation stream (stream
38b), for example. The expanded stream 33a is further cooled in
heat exchanger 12 by heat exchange with cool lean LNG (stream 83b),
cool distillation stream 38a, and demethanizer liquids (stream 39)
as described previously. The further cooled stream 33b enters
separator 13 at -84.degree. F. [-65.degree. C.] and 423 psia [2,916
kPa(a)] where the vapor (stream 34) is separated from the condensed
liquid (stream 35).
[0049] Vapor stream 34 is cooled to -120.degree. F. [-85.degree.
C.] in heat exchanger 14 by heat exchange with cold lean LNG
(stream 83a) and cold distillation stream 38 as described
previously. The partially condensed stream 34a is then supplied to
fractionation tower 20 at a first lower mid-column feed point.
Liquid stream 35 is flash expanded through an appropriate expansion
device, such as expansion valve 17, to the operating pressure of
fractionation tower 20. The expanded stream 35a leaving expansion
valve 17 reaches a temperature of -85.degree. F. [-65.degree. C.]
and is supplied to fractionation tower 20 at a second lower
mid-column feed point.
[0050] The second portion of subcooled stream 81a, reflux stream
36, is expanded to the operating pressure of demethanizer 20 by
expansion valve 15. The expanded stream 36a at -236.degree. F.
[-149.degree. C.] is then supplied as cold top column feed (reflux)
to demethanizer 20. This cold liquid reflux absorbs and condenses
the C.sub.2 components and heavier hydrocarbon components from the
vapors rising in upper rectification section 20a of demethanizer
20.
[0051] The demethanizer in fractionation column 20 is a
conventional distillation column containing a plurality of
vertically spaced trays, one or more packed beds, or some
combination of trays and packing. The fractionation tower 20 may
consist of two sections. The upper absorbing (rectification)
section 20a contains the trays and/or packing to provide the
necessary contact between the vapors rising upward and cold liquid
falling downward to condense and absorb the ethane and heavier
components; the lower stripping (demethanizing) section 20b
contains the trays and/or packing to provide the necessary contact
between the liquids falling downward and the vapors rising upward.
Demethanizing section 20b also includes one or more reboilers (such
as the side reboiler in heat exchanger 12 described previously, and
reboiler 19 using high level utility heat) which heat and vaporize
a portion of the liquids flowing down the column to provide the
stripping vapors which flow up the column. The column liquid stream
40 exits the bottom of the tower at 95.degree. F. [35.degree. C.],
based on a typical specification of a methane to ethane ratio of
0.020:1 on a molar basis in the bottom product, and combines with
stream 80 to form the liquid product (stream 41).
[0052] Overhead distillation stream 38 is withdrawn from the upper
section of fractionation tower 20 at -144.degree. F. [-98.degree.
C.]. It passes countercurrently to the first portion (stream 32a)
of inlet gas stream 31 and vapor stream 34 in heat exchanger 14
where it is heated to -94.degree. F. [-70.degree. C.] (stream 38a),
and countercurrently to the first portion (stream 32) of inlet gas
stream 31 and expanded second portion (stream 33a) in heat
exchanger 12 where it is heated to 13.degree. F. [-11.degree. C.]
(stream 38b). The distillation stream is then re-compressed in two
stages. The first stage is compressor 11 driven by expansion
machine 10. The second stage is compressor 21 driven by a
supplemental power source which compresses stream 38c to sales gas
line pressure (stream 38d). After cooling to 126.degree. F.
[52.degree. C.] in discharge cooler 22, stream 38e combines with
warm lean LNG stream 83c to form the residue gas product (stream
42). Residue gas stream 42 flows to the sales gas pipeline at 1262
psia [8,701 kPa(a)], sufficient to meet line requirements.
[0053] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 3 is set forth in the following
table:
TABLE-US-00003 TABLE III (FIG. 3) Stream Flow Summary - Lb.
Moles/Hr [kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total
31 42,545 5,048 2,972 1,658 53,145 32 5,531 656 386 215 6,909 33
37,014 4,392 2,586 1,443 46,236 34 32,432 1,703 255 29 35,166 35
4,582 2,689 2,331 1,414 11,070 36 7,720 2 0 0 7,773 38 50,165 24 0
0 51,078 40 100 5,026 2,972 1,658 9,840 71 40,293 2,642 491 3
43,689 72/75 4,916 322 60 0 5,330 73/76 35,377 2,320 431 3 38,359
77 35,377 2,320 431 3 38,359 78 0 0 0 0 0 79 45,682 14 0 0 45,990
81 13,162 4 0 0 13,251 83 32,520 10 0 0 32,739 82 5,442 2 0 0 5,478
80 53 2,630 491 3 3,177 42 82,685 34 0 0 83,817 41 153 7,656 3,463
1,661 13,017 Recoveries* Ethane 99.55% Propane 100.00% Butanes+
100.00% Power LNG Feed Pump 3,561 HP [ 5,854 kW] LNG Product Pump
1,740 HP [ 2,861 kW] Residue Gas Compressor 24,852 HP [ 40,856 kW]
Totals 30,153 HP [ 49,571 kW] Low Level Utility Heat Liquid Feed
Heater 65,000 MBTU/Hr [ 41,987 kW] Demethanizer Reboiler 60 19,000
MBTU/Hr [ 12,273 kW] Totals 84,000 MBTU/Hr [ 54,260 kW] High Level
Utility Heat Demethanizer Reboiler 19 41,460 MBTU/Hr [ 26,781 kW]
Demethanizer Reboiler 61 8,400 MBTU/Hr [ 5,426 kW] Totals 49,860
MBTU/Hr [ 32,207 kW] Specific Power HP-Hr/Lb. Mole 2.316 [kW-Hr/kg
mole] [ 3.808 ] *(Based on un-rounded flow rates)
[0054] The improvement offered by the FIG. 3 embodiment of the
present invention is astonishing compared to the FIG. 1 and FIG. 2
processes. Comparing the recovery levels displayed in Table III
above for the FIG. 3 embodiment with those in Table I for the FIG.
1 process shows that the FIG. 3 embodiment of the present invention
improves the ethane recovery from 65.37% to 99.55%, the propane
recovery from 85.83% to 100.00%, and the butanes+ recovery from
99.83% to 100.00%. Further, comparing the utilities consumptions in
Table III with those in Table I shows that although the power
required for the FIG. 3 embodiment of the present invention is
approximately 7% higher than the FIG. 1 process, the process
efficiency of the FIG. 3 embodiment of the present invention is
significantly better than that of the FIG. 1 process. The gain in
process efficiency is clearly seen in the drop in the specific
power, from 2.868 HP-Hr/Lb. Mole [4.715 kW-Hr/kg mole] for the FIG.
1 process to 2.316 HP-Hr/Lb. Mole [3.808 kW-Hr/kg mole] for the
FIG. 3 embodiment of the present invention, an increase of more
than 19% in the production efficiency.
[0055] Comparing the recovery levels displayed in Table III for the
FIG. 3 embodiment with those in Table II for the FIG. 2 processes
shows that the liquids recovery levels are essentially the same.
However, comparing the utilities consumptions in Table III with
those in Table II shows that the power required for the FIG. 3
embodiment of the present invention is about 18% lower than the
FIG. 2 processes. This results in reducing the specific power from
2.851 HP-Hr/Lb. Mole [4.687 kW-Hr/kg mole] for the FIG. 2 processes
to 2.316 HP-Hr/Lb. Mole [3.808 kW-H/kg mole] for the FIG. 3
embodiment of the present invention, an improvement of nearly 19%
in the production efficiency.
[0056] There are six primary factors that account for the improved
efficiency of the present invention. First, compared to many prior
art processes, the present invention does not depend on the LNG
feed itself to directly serve as the reflux for fractionation
column 62. Rather, the refrigeration inherent in the cold LNG is
used in heat exchanger 52 to generate a liquid reflux stream
(stream 82) that contains very little of the C.sub.2 components and
heavier hydrocarbon components that are to be recovered, resulting
in efficient rectification in the absorbing section of
fractionation tower 62 and avoiding the equilibrium limitations of
such prior art processes. Second, splitting the LNG feed into two
portions before feeding fractionation column 62 allows more
efficient use of low level utility heat, thereby reducing the
amount of high level utility heat consumed by reboiler 61. The cold
portion of the LNG feed (stream 75a) serves as a supplemental
reflux stream for fractionation tower 62, providing partial
rectification of the vapors in the expanded vapor and liquid
streams (streams 77a and 78a, respectively) so that heating and at
least partially vaporizing the other portion (stream 73) of the LNG
feed does not unduly increase the condensing load in heat exchanger
52. Third, using a portion of the cold LNG feed (stream 75a) as a
supplemental reflux stream allows using less top reflux (stream
82a) for fractionation tower 62. The lower top reflux flow, plus
the greater degree of heating using low level utility heat in heat
exchanger 53, results in less total liquid feeding fractionation
column 62, reducing the duty required in reboiler 61 and minimizing
the amount of high level utility heat needed to meet the
specification for the bottom liquid product from demethanizer
62.
[0057] Fourth, using the cold lean LNG stream 83a to provide "free"
refrigeration to the gas streams in heat exchangers 12 and 14
eliminates the need for a separate vaporization means (such as heat
exchanger 53 in the FIG. 1 process) to re-vaporize the LNG prior to
delivery to the sales gas pipeline. Fifth, cooling a portion
(stream 32) of inlet gas stream 31 to substantial condensation
prior to expansion to the operating pressure of demethanizer 20
allows the expanded substantially condensed stream 32c to serve as
a supplemental reflux stream for fractionation tower 20, providing
partial rectification of the vapors in the partially condensed
vapor and expanded liquid streams (streams 34a and 35a,
respectively) so that less top reflux (stream 36a) is needed for
fractionation tower 20. Sixth, integrating the LNG plant with the
gas plant allows using a portion (stream 36) of the lean LNG as
reflux for demethanizer 20. The resulting stream 36a is very cold
and contains very little of the C.sub.2 components and heavier
hydrocarbon components that are to be recovered, resulting in very
efficient rectification in absorbing section 20a and further
minimizing the quantity of reflux required for demethanizer 20.
EXAMPLE 2
[0058] An alternative method of processing natural gas is shown in
another embodiment of the present invention as illustrated in FIG.
4. The LNG stream and inlet gas stream compositions and conditions
considered in the process presented in FIG. 4 are the same as those
in FIGS. 1 through 3. Accordingly, the FIG. 4 process can be
compared with the FIGS. 1 and 2 processes to illustrate the
advantages of the present invention, and can likewise be compared
to the embodiment displayed in FIG. 3.
[0059] In the simulation of the FIG. 4 process, the LNG to be
processed (stream 71) from LNG tank 50 enters pump 51 at
-251.degree. F. [-157.degree. C.]. Pump 51 elevates the pressure of
the LNG sufficiently so that it can flow through heat exchangers
and thence to separator 54. Stream 71a exits the pump at
-242.degree. F. [-152.degree. C.] and 1364 psia [9,401 kPa(a)] and
is split into two portions, streams 72 and 73. The first portion,
stream 72, becomes stream 75 and is expanded to the operating
pressure (approximately 415 psia [2,859 kPa(a)]) of fractionation
column 62 by expansion valve 58. The expanded stream 75a leaves
expansion valve 58 at -238.degree. F. [-150C.] and is thereafter
supplied to tower 62 at an upper mid-column feed point.
[0060] The second portion, stream 73, is heated prior to entering
separator 54 so that all or a portion of it is vaporized. In the
example shown in FIG. 4, stream 73 is first heated to -77.degree.
F. [-61 .degree. C.] in heat exchanger 52 by cooling compressed
overhead distillation stream 79a at -70.degree. F. [-57.degree. C.]
and reflux stream 81 at -15.degree. F. [-82.degree. C]. The
partially heated stream 73a becomes stream 76 and is further heated
in heat exchanger 53 using low level utility heat. The heated
stream 76a enters separator 54 at -5.degree. F. [-20.degree. C.]
and 1334 psia [9,195 kPa(a)] where the vapor (stream 77) is
separated from any remaining liquid (stream 78). Vapor stream 77
enters a work expansion machine 55 in which mechanical energy is
extracted from this portion of the high pressure feed. The machine
55 expands the vapor substantially isentropically to the tower
operating pressure, with the work expansion cooling the expanded
stream 77a to a temperature of approximately -107.degree. F.
[-77.degree. C.]. The partially condensed expanded stream 77a is
thereafter supplied as feed to fractionation column 62 at a lower
mid-column feed point. The separator liquid (stream 78), if any, is
expanded to the operating pressure of fractionation column 62 by
expansion valve 59 before expanded stream 78a is supplied to
fractionation tower 62 at a second lower mid-column feed point.
[0061] The column liquid stream 80 exits the bottom of the tower at
54.degree. F. [12.degree. C.], based on a typical specification of
a methane to ethane ratio of 0.020:1 on a molar basis in the bottom
product. Overhead distillation stream 79 is withdrawn from the
upper section of fractionation tower 62 at -144.degree. F.
[-98.degree. C.] and flows to compressor 56 driven by expansion
machine 55, where it is compressed to 805 psia [5,554 kPa(a)]
(stream 79a). At this pressure, the stream is totally condensed as
it is cooled to -115.degree. F. [-82.degree. C.] in heat exchanger
52 as described previously. The condensed liquid (stream 79b) is
then divided into two portions, streams 83 and 81. The first
portion (stream 83) is the methane-rich lean LNG stream, which is
pumped by pump 63 to 1270 psia [8,756 kPa(a)] for subsequent
vaporization in heat exchanger 12, heating stream 83a to 40.degree.
F. [4.degree. C.] as described in paragraph [0063] below to produce
warm lean LNG stream 83b.
[0062] The remaining portion of condensed liquid stream 79b, stream
81, flows to heat exchanger 52 where it is subcooled to
-237.degree. F. [-149.degree. C.] by heat exchange with a portion
of the cold LNG (stream 73) as described previously. The subcooled
stream 81a is then divided into two portions, streams 82 and 36.
The first portion, reflux stream 82, is expanded to the operating
pressure of demethanizer 62 by expansion valve 57. The expanded
stream 82a at -236.degree. F. [-149.degree. C.] is then supplied as
cold top column feed (reflux) to demethanizer 62. This cold liquid
reflux absorbs and condenses the C.sub.2 components and heavier
hydrocarbon components from the vapors rising in the upper
rectification section of demethanizer 62. The disposition of the
second portion, reflux stream 36 for demethanizer 20, is described
in paragraph [0066] below.
[0063] In the simulation of the FIG. 4 process, inlet gas enters
the plant at 126.degree. F. [52.degree. C.] and 600 psia [4,137
kPa(a)] as stream 31. The feed stream 31 is divided into two
portions, streams 32 and 33. The first portion, stream 32, is
cooled in heat exchanger 12 by heat exchange with cold lean LNG
(stream 83a) at -96.degree. F. [-71.degree. C.], cool compressed
distillation stream 38b at -109.degree. F. [-78.degree. C.], and
demethanizer liquids (stream 39) at -63.degree. F. [-53.degree.
C.]. The partially cooled stream 32a is further cooled from
-96.degree. F. [-71.degree. C.] to -121.degree. F. [-85.degree. C.]
in heat exchanger 14 by heat exchange with cold compressed
distillation stream 38a at -128.degree. F. [-89.degree. C.]. The
substantially condensed stream 32b is then flash expanded through
an appropriate expansion device, such as expansion valve 16, to the
operating pressure (approximately 443 psia [3,052 kPa(a)]) of
fractionation tower 20, cooling stream 32c to -129.degree. F.
[-90.degree. C.] before it is supplied to fractionation tower 20 at
an upper mid-column feed point.
[0064] The second portion of feed stream 31, stream 33, is cooled
in heat exchanger 12 by heat exchange with cold lean LNG (stream
83a), cool compressed distillation stream 38b, and demethanizer
liquids (stream 39) as described previously. The cooled stream 33a
enters separator 13 at -86.degree. F. [-65.degree. C.] and 584 psia
[4,027 kPa(a)] where the vapor (stream 34) is separated from the
condensed liquid (stream 35). Liquid stream 35 is flash expanded
through an appropriate expansion device, such as expansion valve
17, to the operating pressure of fractionation tower 20. The
expanded stream 35a leaving expansion valve 17 reaches a
temperature of -100.degree. F. [-73.degree. C.] and is supplied to
fractionation tower 20 at a first lower mid-column feed point.
[0065] The vapor from separator 13 (stream 34) enters a work
expansion machine 10 in which mechanical energy is extracted from
this portion of the high pressure feed. The machine 10 expands the
vapor substantially isentropically to slightly above the tower
operating pressure, with the work expansion cooling the expanded
stream 34a to a temperature of approximately -106.degree. F.
[-77.degree. C.]. The expanded stream 34a is further cooled to
-121.degree. F. [-85.degree. C.] in heat exchanger 14 by heat
exchange with cold compressed distillation stream 38a as described
previously, whereupon the partially condensed expanded stream 34b
is thereafter supplied to fractionation tower 20 at a second lower
mid-column feed point.
[0066] The second portion of subcooled stream 81a, reflux stream
36, is expanded to the operating pressure of demethanizer 20 by
expansion valve 15. The expanded stream 36a at -236.degree. F.
[-149.degree. C.] is then supplied as cold top column feed (reflux)
to demethanizer 20. This cold liquid reflux absorbs and condenses
the C.sub.2 components and heavier hydrocarbon components from the
vapors rising in the upper rectification section of demethanizer
20.
[0067] The column liquid stream 40 exits the bottom of the tower at
102.degree. F. [39.degree. C.], based on a typical specification of
a methane to ethane ratio of 0.020:1 on a molar basis in the bottom
product, and combines with stream 80 to form the liquid product
(stream 41). Overhead distillation stream 38 is withdrawn from the
upper section of fractionation tower 20 at -141.degree. F.
[-96.degree. C.] and flows to compressor 11 driven by expansion
machine 10, where it is compressed to 501 psia [3,452 kPa(a)]. The
cold compressed distillation stream 38a passes countercurrently to
the first portion (stream 32a) of inlet gas stream 31 and expanded
vapor stream 34a in heat exchanger 14 where it is heated to
-109.degree. F. [-78.degree. C.] (stream 38b), and countercurrently
to the first portion (stream 32) and second portion (stream 33) of
inlet gas stream 31 in heat exchanger 12 where it is heated to
31.degree. F. [-1.degree. C.] (stream 38c). The heated distillation
stream then enters compressor 21 driven by a supplemental power
source which compresses stream 38c to sales line pressure (stream
38d). After cooling to 126.degree. F. [52.degree. C.] in discharge
cooler 22, stream 38e combines with warm lean LNG stream 83b to
form the residue gas product (stream 42). Residue gas stream 42
flows to the sales gas pipeline at 1262 psia [8,701 kPa(a)],
sufficient to meet line requirements.
[0068] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 4 is set forth in the following
table:
TABLE-US-00004 TABLE IV (FIG. 4) Stream Flow Summary - Lb. Moles/Hr
[kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31
42,545 5,048 2,972 1,658 53,145 32 3,404 404 238 133 4,251 33
39,141 4,644 2,734 1,525 48,894 34 28,606 1,181 191 26 30,730 35
10,535 3,463 2,543 1,499 18,164 36 8,046 2 0 0 8,101 38 50,491 27 0
0 51,413 40 100 5,023 2,972 1,658 9,833 71 40,293 2,642 491 3
43,689 72/75 4,916 322 60 0 5,330 73/76 35,377 2,320 431 3 38,359
77 35,377 2,320 431 3 38,359 78 0 0 0 0 0 79 45,682 14 0 0 45,990
81 13,488 4 0 0 13,579 83 32,194 10 0 0 32,411 82 5,442 2 0 0 5,478
80 53 2,630 491 3 3,177 42 82,685 37 0 0 83,824 41 153 7,653 3,463
1,661 13,010 Recoveries* Ethane 99.51% Propane 100.00% Butanes+
100.00% Power LNG Feed Pump 3,561 HP [ 5,854 kW] LNG Product Pump
1,727 HP [ 2,839 kW] Residue Gas Compressor 24,400 HP [ 40,113 kW]
Totals 29,688 HP [ 48,806 kW] Low Level Utility Heat Liquid Feed
Heater 65,000 MBTU/Hr [ 41,987 kW] Demethanizer Reboiler 60 19,000
MBTU/Hr [ 12,273 kW] Totals 84,000 MBTU/Hr [ 54,260 kW] High Level
Utility Heat Demethanizer Reboiler 19 37,360 MBTU/Hr [ 24,133 kW]
Demethanizer Reboiler 61 8,400 MBTU/Hr [ 5,426 kW] Totals 45,760
MBTU/Hr [ 29,559 kW] Specific Power HP-Hr/Lb. Mole 2.282 [kW-Hr/kg
mole] [ 3.751 ] *(Based on un-rounded flow rates)
[0069] A comparison of Tables III and IV shows that the FIG. 4
embodiment of the present invention achieves essentially the same
liquids recovery as the FIG. 3 embodiment. However, the FIG. 4
embodiment uses less power than the FIG. 3 embodiment, improving
the specific power by slightly more than 1%. In addition, the high
level utility heat required for the FIG. 4 embodiment of the
present invention is about 8% less than that of the FIG. 3
embodiment.
EXAMPLE 3
[0070] Another alternative method of processing natural gas is
shown in the embodiment of the present invention as illustrated in
FIG. 5. The LNG stream and inlet gas stream compositions and
conditions considered in the process presented in FIG. 5 are the
same as those in FIGS. 1 through 4. Accordingly, the FIG. 5 process
can be compared with the FIGS. 1 and 2 processes to illustrate the
advantages of the present invention, and can likewise be compared
to the embodiments displayed in FIGS. 3 and 4.
[0071] In the simulation of the FIG. 5 process, the LNG to be
processed (stream 71) from LNG tank 50 enters pump 51 at
-251.degree. F. [-157.degree. C.]. Pump 51 elevates the pressure of
the LNG sufficiently so that it can flow through heat exchangers
and thence to separator 54. Stream 71a exits the pump at
-242.degree. F. [-152.degree. C.] and 1364 psia [9,401 kPa(a)] and
is split into two portions, streams 72 and 73. The first portion,
stream 72, becomes stream 75 and is expanded to the operating
pressure (approximately 415 psia [2,859 kPa(a)]) of fractionation
column 62 by expansion valve 58. The expanded stream 75a leaves
expansion valve 58 at -238.degree. F. [-150C] and is thereafter
supplied to tower 62 at an upper mid-column feed point.
[0072] The second portion, stream 73, is heated prior to entering
separator 54 so that all or a portion of it is vaporized. In the
example shown in FIG. 5, stream 73 is first heated to -77.degree.
F. [-61.degree. C.] in heat exchanger 52 by cooling compressed
overhead distillation stream 79a at -70.degree. F. [-57.degree. C.]
and reflux stream 81 at -12.degree. F. [-80.degree. C]. The
partially heated stream 73a becomes stream 76 and is further heated
in heat exchanger 53 using low level utility heat. The heated
stream 76a enters separator 54 at -5.degree. F. [-20.degree. C.]
and 1334 psia [9,195 kPa(a)] where the vapor (stream 77) is
separated from any remaining liquid (stream 78). Vapor stream 77
enters a work expansion machine 55 in which mechanical energy is
extracted from this portion of the high pressure feed. The machine
55 expands the vapor substantially isentropically to the tower
operating pressure, with the work expansion cooling the expanded
stream 77a to a temperature of approximately -107.degree. F.
[-77.degree. C.]. The partially condensed expanded stream 77a is
thereafter supplied as feed to fractionation column 62 at a lower
mid-column feed point. The separator liquid (stream 78), if any, is
expanded to the operating pressure of fractionation column 62 by
expansion valve 59 before expanded stream 78a is supplied to
fractionation tower 62 at a second lower mid-column feed point.
[0073] The column liquid stream 80 exits the bottom of the tower at
54.degree. F. [12.degree. C.], based on a typical specification of
a methane to ethane ratio of 0.020:1 on a molar basis in the bottom
product. Overhead distillation stream 79 is withdrawn from the
upper section of fractionation tower 62 at -144.degree. F.
[-98.degree. C.] and flows to compressor 56 driven by expansion
machine 55, where it is compressed to 805 psia [5,554 kPa(a)]
(stream 79a). At this pressure, the stream is totally condensed as
it is cooled to -112.degree. F. [-80.degree. C.] in heat exchanger
52 as described previously. The condensed liquid (stream 79b) is
then divided into two portions, streams 83 and 81. The first
portion (stream 83) is the methane-rich lean LNG stream, which is
pumped by pump 63 to 1270 psia [8,756 kPa(a)] for subsequent
vaporization in heat exchanger 12, heating stream 83a to 40.degree.
F. [4.degree. C.] as described in paragraph [0075] below to produce
warm lean LNG stream 83b.
[0074] The remaining portion of condensed liquid stream 79b, stream
81, flows to heat exchanger 52 where it is subcooled to
-237.degree. F. [-149.degree. C.] by heat exchange with a portion
of the cold LNG (stream 73) as described previously. The subcooled
stream 81a is then divided into two portions, streams 82 and 36.
The first portion, reflux stream 82, is expanded to the operating
pressure of demethanizer 62 by expansion valve 57. The expanded
stream 82a at -236.degree. F. [-149.degree. C.] is then supplied as
cold top column feed (reflux) to demethanizer 62. This cold liquid
reflux absorbs and condenses the C.sub.2 components and heavier
hydrocarbon components from the vapors rising in the upper
rectification section of demethanizer 62. The disposition of the
second portion, reflux stream 36 for demethanizer 20, is described
in paragraph [0078] below.
[0075] In the simulation of the FIG. 5 process, inlet gas enters
the plant at 126.degree. F. [52.degree. C.] and 600 psia [4,137
kPa(a)] as stream 31. The feed stream 31 is divided into two
portions, streams 32 and 33. The first portion, stream 32, is
cooled in heat exchanger 12 by heat exchange with cold lean LNG
(stream 83a) at -89.degree. F. [-67.degree. C.], cool compressed
distillation stream 38b at -91.degree. F. [-68.degree. C.], and
demethanizer liquids (stream 39) at -89.degree. F. [-67.degree.
C.]. The partially cooled stream 32a is further cooled from
-86.degree. F. [-65.degree. C.] to -100.degree. F. [-74.degree. C.]
in heat exchanger 14 by heat exchange with cold compressed
distillation stream 38a at -112.degree. F. [-80.degree. C.]. The
substantially condensed stream 32b is then flash expanded through
an appropriate expansion device, such as expansion valve 16, to the
operating pressure (approximately 428 psia [2,949 kPa(a)]) of
fractionation tower 20, cooling stream 32c to -117.degree. F.
[-83.degree. C.] before it is supplied to fractionation tower 20 at
an upper mid-column feed point.
[0076] The second portion of feed stream 31, stream 33, enters a
work expansion machine 10 in which mechanical energy is extracted
from this portion of the high pressure feed. The machine 10 expands
the vapor substantially isentropically to a pressure slightly above
the operating pressure of fractionation tower 20, with the work
expansion cooling the expanded stream 33a to a temperature of
approximately 95.degree. F. [35.degree. C.]. The expanded stream
33a is further cooled in heat exchanger 12 by heat exchange with
cold lean LNG (stream 83a), cool compressed distillation stream
38b, and demethanizer liquids (stream 39) as described previously.
The further cooled stream 33b enters separator 13 at -85.degree. F.
[-65.degree. C.] and 436 psia [3,004 kPa(a)] where the vapor
(stream 34) is separated from the condensed liquid (stream 35).
[0077] Vapor stream 34 is cooled to -100.degree. F. [-74.degree.
C.] in heat exchanger 14 by heat exchange with cold compressed
distillation stream 38a as described previously. The partially
condensed stream 34a is then supplied to fractionation tower 20 at
a first lower mid-column feed point. Liquid stream 35 is flash
expanded through an appropriate expansion device, such as expansion
valve 17, to the operating pressure of fractionation tower 20. The
expanded stream 35a leaving expansion valve 17 reaches a
temperature of -86.degree. F. [-65.degree. C.] and is supplied to
fractionation tower 20 at a second lower mid-column feed point.
[0078] The second portion of subcooled stream 81a, reflux stream
36, is expanded to the operating pressure of demethanizer 20 by
expansion valve 15. The expanded stream 36a at -236.degree. F.
[-149.degree. C.] is then supplied as cold top column feed (reflux)
to demethanizer 20. This cold liquid reflux absorbs and condenses
the C.sub.2 components and heavier hydrocarbon components from the
vapors rising in the upper rectification section of demethanizer
20.
[0079] The column liquid stream 40 exits the bottom of the tower at
98.degree. F. [37.degree. C.], based on a typical specification of
a methane to ethane ratio of 0.020:1 on a molar basis in the bottom
product, and combines with stream 80 to form the liquid product
(stream 41). Overhead distillation stream 38 is withdrawn from the
upper section of fractionation tower 20 at -143.degree. F.
[-97.degree. C.] and flows to compressor 11 driven by expansion
machine 10, where it is compressed to 573 psia [3,950 kPa(a)]. The
cold compressed distillation stream 38a passes countercurrently to
the first portion (stream 32a) of inlet gas stream 31 and vapor
stream 34 in heat exchanger 14 where it is heated to -91.degree. F.
[-68.degree. C.] (stream 38b), and countercurrently to the first
portion (stream 32) and expanded second portion (stream 33a) of
inlet gas stream 31 in heat exchanger 12 where it is heated to
67.degree. F. [19.degree. C.] (stream 38c). The heated distillation
stream then enters compressor 21 driven by a supplemental power
source which compresses stream 38c to sales line pressure (stream
38d). After cooling to 126.degree. F. [52.degree. C.] in discharge
cooler 22, stream 38e combines with warm lean LNG stream 83b to
form the residue gas product (stream 42). Residue gas stream 42
flows to the sales gas pipeline at 1262 psia [8,701 kPa(a)],
sufficient to meet line requirements.
[0080] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 5 is set forth in the following
table:
TABLE-US-00005 TABLE V (FIG. 5) Stream Flow Summary - Lb. Moles/Hr
[kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31
42,545 5,048 2,972 1,658 53,145 32 14,465 1,716 1,010 564 18,069 33
28,080 3,332 1,962 1,094 35,076 34 24,317 1,236 184 21 26,322 35
3,763 2,096 1,778 1,073 8,754 36 10,372 3 0 0 10,442 38 52,817 30 0
0 53,749 40 100 5,021 2,972 1,658 9,838 71 40,293 2,642 491 3
43,689 72/75 4,916 322 60 0 5,330 73/76 35,377 2,320 431 3 38,359
77 35,377 2,320 431 3 38,359 78 0 0 0 0 0 79 45,682 14 0 0 45,990
81 15,814 5 0 0 15,920 83 29,868 9 0 0 30,070 82 5,442 2 0 0 5,478
80 53 2,630 491 3 3,177 42 82,685 39 0 0 83,819 41 153 7,651 3,463
1,661 13,015 Recoveries* Ethane 99.48% Propane 100.00% Butanes+
100.00% Power LNG Feed Pump 3,561 HP [ 5,854 kW] LNG Product Pump
1,778 HP [ 2,923 kW] Residue Gas Compressor 23,201 HP [ 38,142 kW]
Totals 28,540 HP [ 46,919 kW] Low Level Utility Heat Liquid Feed
Heater 65,000 MBTU/Hr [ 41,987 kW] Demethanizer Reboiler 60 19,000
MBTU/Hr [ 12,273 kW] Totals 84,000 MBTU/Hr [ 54,260 kW] High Level
Utility Heat Demethanizer Reboiler 19 53,370 MBTU/Hr [ 34,475 kW]
Demethanizer Reboiler 61 8,400 MBTU/Hr [ 5,426 kW] Totals 61,770
MBTU/Hr [ 39,901 kW] Specific Power HP-Hr/Lb. Mole 2.193 [kW-Hr/kg
mole] [ 3.605 ] *(Based on un-rounded flow rates)
[0081] A comparison of Tables III, IV, and V shows that the FIG. 5
embodiment of the present invention achieves essentially the same
liquids recovery as the FIG. 3 and FIG. 4 embodiments. The FIG. 5
embodiment uses less power than the FIG. 3 and FIG. 4 embodiments,
improving the specific power by over 5% relative to the FIG. 3
embodiment and nearly 4% relative to the FIG. 4 embodiment.
However, the high level utility heat required for the FIG. 5
embodiment of the present invention is somewhat higher than that of
the FIG. 3 and FIG. 4 embodiments (by 24% and 35%, respectively).
The choice of which embodiment to use for a particular application
will generally be dictated by the relative costs of power and high
level utility heat and the relative capital costs of pumps, heat
exchangers, and compressors.
EXAMPLE 4
[0082] An alternative method of processing LNG and natural gas is
shown in the embodiment of the present invention as illustrated in
FIG. 6. The LNG stream and inlet gas stream compositions and
conditions considered in the process presented in FIG. 6 are the
same as those in FIGS. 1 through 5. Accordingly, the FIG. 5 process
can be compared with the FIGS. 1 and 2 processes to illustrate the
advantages of the present invention, and can likewise be compared
to the embodiments displayed in FIGS. 3 through 5.
[0083] In the simulation of the FIG. 6 process, the LNG to be
processed (stream 71) from LNG tank 50 enters pump 51 at
-251.degree. F. [-157.degree. C.]. Pump 51 elevates the pressure of
the LNG sufficiently so that it can flow through heat exchangers
and thence to separator 54. Stream 71a exits the pump at
-242.degree. F. [-152.degree. C.] and 1364 psia [9,401 kPa(a)] and
is split into two portions, streams 72 and 73. The first portion,
stream 72, becomes stream 75 and is expanded to the operating
pressure (approximately 435 psia [2,997 kPa(a)]) of fractionation
column 20 by expansion valve 58. The expanded stream 75a leaves
expansion valve 58 at -238.degree. F. [-150.degree. C.] and is
thereafter supplied to tower 20 at a first upper mid-column feed
point.
[0084] The second portion, stream 73, is heated prior to entering
separator 54 so that all or a portion of it is vaporized. In the
example shown in FIG. 6, stream 73 is first heated to -76.degree.
F. [-60.degree. C.] in heat exchanger 52 by cooling compressed
overhead distillation stream 81a at -65.degree. F. [-54.degree. C.]
and reflux stream 82 at -117.degree. F. [-82.degree. C.], exchanger
14 as described in paragraph [0085] below. The partially heated
stream 73b becomes stream 76 and is further heated in heat
exchanger 53 using low level utility heat. The heated stream 76a
enters separator 54 at -5.degree. F. [-20.degree. C.] and 1334 psia
[9,195 kPa(a)] where the vapor (stream 77) is separated from any
remaining liquid (stream 78). Vapor stream 77 enters a work
expansion machine 55 in which mechanical energy is extracted from
this portion of the high pressure feed. The machine 55 expands the
vapor substantially isentropically to the tower operating pressure,
with the work expansion cooling the expanded stream 77a to a
temperature of approximately -104.degree. F. [-76.degree. C.]. The
partially condensed expanded stream 77a is thereafter supplied as
feed to fractionation column 20 at a first lower mid-column feed
point. The separator liquid (stream 78), if any, is expanded to the
operating pressure of fractionation column 20 by expansion valve 59
before expanded stream 78a is supplied to fractionation tower 20 at
a second lower mid-column feed point.
[0085] In the simulation of the FIG. 6 process, inlet gas enters
the plant at 126.degree. F. [52.degree. C.] and 600 psia [4,137
kPa(a)] as stream 31. The feed stream 31 is divided into two
portions, streams 32 and 33. The first portion, stream 32, is
cooled in heat exchanger 12 by heat exchange with cold lean LNG
(stream 83a) at -103.degree. F. [-75.degree. C.], cool compressed
distillation stream 38b at -92.degree. F. [-69.degree. C.], and
demethanizer liquids (stream 39) at -78.degree. F. [-61.degree.
C.]. The partially cooled stream 32a is further cooled from
-94.degree. F. [-70.degree. C.] to -101.degree. F. [-74.degree. C.]
in heat exchanger 14 by heat exchange with the partially heated
second portion (stream 73a) of the LNG stream and with cold
compressed distillation stream 38a at -106.degree. F. [-77.degree.
C.]. The substantially condensed stream 32b is then flash expanded
through an appropriate expansion device, such as expansion valve
16, to the operating pressure of fractionation tower 20, cooling
stream 32c to -117.degree. F. [-83.degree. C.] before it is
supplied to fractionation tower 20 at a second upper mid-column
feed point.
[0086] The second portion of feed stream 31, stream 33, enters a
work expansion machine 10 in which mechanical energy is extracted
from this portion of the high pressure feed. The machine 10 expands
the vapor substantially isentropically to a pressure slightly above
the operating pressure of fractionation tower 20, with the work
expansion cooling the expanded stream 33a to a temperature of
approximately 96.degree. F. [36.degree. C.]. The expanded stream
33a is further cooled in heat exchanger 12 by heat exchange with
cold lean LNG (stream 83a), cool compressed distillation stream
38b, and demethanizer liquids (stream 39) as described previously.
The further cooled stream 33b enters separator 13 at -90.degree. F.
[-68.degree. C.] and 443 psia [3,052 kPa(a)] where the vapor
(stream 34) is separated from the condensed liquid (stream 35).
[0087] Vapor stream 34 is cooled to -101.degree. F. [-74.degree.
C.] in heat exchanger 14 by heat exchange with the partially heated
second portion (stream 73a) of the LNG stream and with cold
compressed distillation stream 38a as described previously. The
partially condensed stream 34a is then supplied to fractionation
tower 20 at a third lower mid-column feed point. Liquid stream 35
is flash expanded through an appropriate expansion device, such as
expansion valve 17, to the operating pressure of fractionation
tower 20. The expanded stream 35a leaving expansion valve 17
reaches a temperature of -90.degree. F. [-68.degree. C.] and is
supplied to fractionation tower 20 at a fourth lower mid-column
feed point.
[0088] The liquid product stream 41 exits the bottom of the tower
at 89.degree. F. [32.degree. C.], based on a typical specification
of a methane to ethane ratio of 0.020:1 on a molar basis in the
bottom product. Overhead distillation stream 79 is withdrawn from
the upper section of fractionation tower 20 at -142.degree. F.
[-97.degree. C.] and is divided into two portions, stream 81 and
stream 38. The first portion (stream 81) flows to compressor 56
driven by expansion machine 55, where it is compressed to 864 psia
[5,955 kPa(a)] (stream 81a). At this pressure, the stream is
totally condensed as it is cooled to -117.degree. F. [-83.degree.
C.] in heat exchanger 52 as described previously. The condensed
liquid (stream 81b) is then divided into two portions, streams 83
and 82. The first portion (stream 83) is the methane-rich lean LNG
stream, which is pumped by pump 63 to 1270 psia [8,756 kPa(a)] for
subsequent vaporization in heat exchanger 12, heating stream 83a to
40.degree. F. [4.degree. C.] as described previously to produce
warm lean LNG stream 83b.
[0089] The remaining portion of stream 81b (stream 82) flows to
heat exchanger 52 where it is subcooled to -237.degree. F.
[-149.degree. C.] by heat exchange with a portion of the cold LNG
(stream 73) as described previously. The subcooled stream 82a is
expanded to the operating pressure of fractionation column 20 by
expansion valve 57. The expanded stream 82b at -236.degree. F.
[-149.degree. C.] is then supplied as cold top column feed (reflux)
to demethanizer 20. This cold liquid reflux absorbs and condenses
the C.sub.2 components and heavier hydrocarbon components from the
vapors rising in the upper rectification section of demethanizer
20.
[0090] The second portion of distillation stream 79 (stream 38)
flows to compressor 11 driven by expansion machine 10, where it is
compressed to 604 psia [4,165 kPa(a)]. The cold compressed
distillation stream 38a passes countercurrently to the first
portion (stream 32a) of inlet gas stream 31 and vapor stream 34 in
heat exchanger 14 where it is heated to -92.degree. F. [-69.degree.
C.] (stream 38b), and countercurrently to the first portion (stream
32) and expanded second portion (stream 33a) of inlet gas stream 31
in heat exchanger 12 where it is heated to 48.degree. F. [9.degree.
C.] (stream 38c). The heated distillation stream then enters
compressor 21 driven by a supplemental power source which
compresses stream 38c to sales line pressure (stream 38d). After
cooling to 126.degree. F. [52.degree. C.] in discharge cooler 22,
stream 38e combines with warm lean LNG stream 83b to form the
residue gas product (stream 42). Residue gas stream 42 flows to the
sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient to meet
line requirements.
[0091] A summary of stream flow rates and energy consumption for
the process illustrated in FIG. 6 is set forth in the following
table:
TABLE-US-00006 TABLE VI (FIG. 6) Stream Flow Summary - Lb. Moles/Hr
[kg moles/Hr] Stream Methane Ethane Propane Butanes+ Total 31
42,545 5,048 2,972 1,658 53,145 32 7,871 934 550 307 9,832 33
34,674 4,114 2,422 1,351 43,313 34 29,159 1,328 185 21 31,380 35
5,515 2,786 2,237 1,330 11,933 71 40,293 2,642 491 3 43,689 72/75
5,037 330 61 0 5,461 73/76 35,256 2,312 430 3 38,228 77 35,256
2,312 430 3 38,228 78 0 0 0 0 0 79 97,329 46 0 0 98,696 38 54,991
26 0 0 55,763 81 42,338 20 0 0 42,933 82 14,644 7 0 0 14,850 83
27,694 13 0 0 28,083 42 82,685 39 0 0 83,846 41 153 7,651 3,463
1,661 12,988 Recoveries* Ethane 99.48% Propane 100.00% Butanes+
100.00% Power LNG Feed Pump 3,561 HP [ 5,854 kW] LNG Product Pump
1,216 HP [ 1,999 kW] Residue Gas Compressor 21,186 HP [ 34,829 kW]
Totals 25,963 HP [ 42,682 kW] Low Level Utility Heat Liquid Feed
Heater 70,000 MBTU/Hr [ 45,217 kW] Demethanizer Reboiler 18 30,000
MBTU/Hr [ 19,378 kW] Totals 100,000 MBTU/Hr [ 64,595 kW] High Level
Utility Heat Demethanizer Reboiler 19 39,180 MBTU/Hr [ 25,308 kW]
Specific Power HP-Hr/Lb. Mole 1.999 [kW-Hr/kg mole] [ 3.286 ]
*(Based on un-rounded flow rates)
[0092] A comparison of Tables III, IV, V, and VI shows that the
FIG. 6 embodiment of the present invention achieves essentially the
same liquids recovery as the FIGS. 3, 4, and 5 embodiments.
However, the reduction in the energy consumption of the FIG. 6
embodiment of the present invention relative to the embodiments in
FIGS. 3 through 5 is unexpectedly large. The FIG. 6 embodiment uses
less power than the FIGS. 3, 4, and 5 embodiments, reducing the
specific power by 14%, 12%, and 9%, respectively. The high level
utility heat required for the FIG. 6 embodiment of the present
invention is also lower than that of the FIGS. 3, 4, and 5
embodiments (by 21%, 14%, and 37%, respectively). These large gains
in process efficiency are mainly due to the more optimal
distribution of the column feeds afforded by integrating the LNG
processing and the natural gas processing into a single
fractionation column, demethanizer 20. For instance, the relative
distribution of the inlet gas stream 31 between stream 32 (which
forms the substantially condensed expanded stream 32c) and stream
33 supplied to expansion machine 10 can be optimized for power
production, since stream 75a from LNG stream 71 provides part of
the supplemental rectification for column 20 that must be provided
entirely by stream 32c in the FIGS. 3 through 5 embodiments.
[0093] The capital cost of the FIG. 6 embodiment of the present
invention will generally be less than that of the FIGS. 3, 4, and 5
embodiments since it uses only one fractionation column, and due to
the reduction in power and high level utility heat consumption. The
choice of which embodiment to use for a particular application will
generally be dictated by the relative costs of power and high level
utility heat and the relative capital costs of columns, pumps, heat
exchangers, and compressors.
Other Embodiments
[0094] Some circumstances may favor using cold distillation stream
38 in the FIG. 6 embodiment for heat exchange prior to compression
as shown in the embodiment displayed in FIG. 7. In other instances,
work expansion of the high pressure inlet gas may be more
advantageous after cooling and separation of any liquids, as shown
in the embodiment displayed in FIG. 8. The choices regarding the
streams used for work expansion and where best to apply the power
generated in compressing the process streams will depend on such
factors as inlet gas pressure and composition, and must be
determined for each application.
[0095] When the inlet gas is leaner, separator 13 in FIGS. 3
through 8 may not be needed. Depending on the quantity of heavier
hydrocarbons in the feed gas and the feed gas pressure, the cooled
stream 33b (FIGS. 3, 5, 6, and 7) or cooled stream 33a (FIGS. 4 and
8) leaving heat exchanger 12 may not contain any liquid (because it
is above its dewpoint, or because it is above its cricondenbar), so
that separator 13 may not be justified. In such cases, separator 13
and expansion valve 17 may be eliminated as shown by the dashed
lines. When the LNG to be processed is lean or when complete
vaporization of the LNG in heat exchangers 52 and 53 is
contemplated, separator 54 in FIGS. 3 through 8 may not be
justified. Depending on the quantity of heavier hydrocarbons in the
inlet LNG and the pressure of the LNG stream leaving feed pump 51,
the heated LNG stream leaving heat exchanger 53 may not contain any
liquid (because it is above its dewpoint, or because it is above
its cricondenbar). In such cases, separator 54 and expansion valve
59 may be eliminated as shown by the dashed lines.
[0096] In the embodiments of the present invention illustrated in
FIGS. 4 and 8, the expanded substantially condensed stream 32c is
formed using a portion (stream 32) of inlet gas stream 31.
Depending on the feed gas composition and other factors, some
circumstances may favor using a portion of the vapor (stream 34)
from separator 13 instead. In such instances, a portion of the
separator 13 vapor forms stream 32a as shown by the dashed lines in
FIGS. 4 and 8, with the remaining portion forming the stream 34
that is fed to expansion machine 10.
[0097] In the examples shown, total condensation of stream 79b in
FIGS. 3 through 5 and stream 81b in FIGS. 6 through 8 is shown.
Some circumstances may favor subcooling these streams, while other
circumstances may favor only partial condensation. Should partial
condensation of these streams be achieved, processing of the
uncondensed vapor may be necessary, using a compressor or other
means to elevate the pressure of the vapor so that it can join the
pumped condensed liquid. Alternatively, the uncondensed vapor could
be routed to the plant fuel system or other such use.
[0098] Feed gas conditions, LNG conditions, plant size, available
equipment, or other factors may indicate that elimination of work
expansion machines 10 and/or 55, or replacement with an alternate
expansion device (such as an expansion valve), is feasible.
Although individual stream expansion is depicted in particular
expansion devices, alternative expansion means may be employed
where appropriate.
[0099] In FIGS. 3 through 8, individual heat exchangers have been
shown for most services. However, it is possible to combine two or
more heat exchange services into a common heat exchanger, such as
combining heat exchangers 12 and 14 in FIGS. 3 through 8 into a
common heat exchanger. In some cases, circumstances may favor
splitting a heat exchange service into multiple exchangers. The
decision as to whether to combine heat exchange services or to use
more than one heat exchanger for the indicated service will depend
on a number of factors including, but not limited to, inlet gas
flow rate, LNG flow rate, heat exchanger size, stream temperatures,
etc. In accordance with the present invention, the use and
distribution of the methane-rich lean LNG and tower overhead
streams for process heat exchange, and the particular arrangement
of heat exchangers for heating the LNG streams and cooling the feed
gas streams, must be evaluated for each particular application, as
well as the choice of process streams for specific heat exchange
services.
[0100] In the embodiments of the present invention illustrated in
FIGS. 3 through 8, lean LNG stream 83a is used directly to provide
cooling in heat exchanger 12 or heat exchangers 12 and 14. However,
some circumstances may favor using the lean LNG to cool an
intermediate heat transfer fluid, such as propane or other suitable
fluid, whereupon the cooled heat transfer fluid is then used to
provide cooling in heat exchanger 12 or heat exchangers 12 and 14.
This alternative means of indirectly using the refrigeration
available in lean LNG stream 83a accomplishes the same process
objectives as the direct use of stream 83a for cooling in the FIGS.
3 through 8 embodiments of the present invention. The choice of how
best to use the lean LNG stream for refrigeration will depend
mainly on the composition of the inlet gas, but other factors may
affect the choice as well.
[0101] It will be recognized that the relative amount of feed found
in each branch of the split LNG feed to fractionation column 62, in
each branch of the split inlet gas to fractionation column 20, and
in each branch of the split LNG feed and the split inlet gas to
fractionation column 20 will depend on several factors, including
inlet gas composition, LNG composition, the amount of heat which
can economically be extracted from the feed, and the quantity of
horsepower available. More feed to the top of the column may
increase recovery while increasing the duty in reboilers 61 and/or
19 and thereby increasing the high level utility heat requirements.
Increasing feed lower in the column reduces the high level utility
heat consumption but may also reduce product recovery. The relative
locations of the mid-column feeds may vary depending on inlet gas
composition, LNG composition, or other factors such as the desired
recovery level and the amount of vapor formed during heating of the
LNG streams. Moreover, two or more of the feed streams, or portions
thereof, may be combined depending on the relative temperatures and
quantities of individual streams, and the combined stream then fed
to a mid-column feed position.
[0102] In some circumstance it may be desirable to recover
refrigeration from the portion (stream 75a) of LNG feed stream 71
that is fed to an upper mid-column feed point on demethanizer 62
(FIGS. 3 through 5) and demethanizer 20 (FIGS. 6 through 8). In
such cases, all of stream 71a would be directed to heat exchanger
52 (stream 73) and the partially heated LNG stream (stream 73a in
FIGS. 3 through 5 and stream 73b in FIGS. 6 through 8) would then
be divided into stream 76 and stream 74 (as shown by the dashed
lines), whereupon stream 74 would be directed to stream 75.
[0103] In the examples given for the FIGS. 3 through 6 embodiments,
recovery of C.sub.2 components and heavier hydrocarbon components
is illustrated. However, it is believed that the FIGS. 3 through 8
embodiments are also advantageous when recovery of only C.sub.3
components and heavier hydrocarbon components is desired. The
present invention provides improved recovery of C.sub.2 components
and heavier hydrocarbon components or of C.sub.3 components and
heavier hydrocarbon components per amount of utility consumption
required to operate the process. An improvement in utility
consumption required for operating the process may appear in the
form of reduced power requirements for compression or pumping,
reduced energy requirements for tower reboilers, or a combination
thereof. Alternatively, the advantages of the present invention may
be realized by accomplishing higher recovery levels for a given
amount of utility consumption, or through some combination of
higher recovery and improvement in utility consumption.
[0104] While there have been described what are believed to be
preferred embodiments of the invention, those skilled in the art
will recognize that other and further modifications may be made
thereto, e.g. to adapt the invention to various conditions, types
of feed, or other requirements without departing from the spirit of
the present invention as defined by the following claims.
* * * * *