U.S. patent application number 12/466661 was filed with the patent office on 2010-11-18 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 | 20100287982 12/466661 |
Document ID | / |
Family ID | 43067387 |
Filed Date | 2010-11-18 |
United States Patent
Application |
20100287982 |
Kind Code |
A1 |
Martinez; Tony L. ; et
al. |
November 18, 2010 |
Liquefied Natural Gas and Hydrocarbon Gas Processing
Abstract
A process for the recovery of heavier hydrocarbons from a
liquefied natural gas (LNG) stream and a hydrocarbon gas stream is
disclosed. The LNG feed stream is heated to vaporize at least part
of it, then expanded and supplied to a fractionation column at a
first mid-column feed position. The gas stream is expanded and
cooled, then supplied to the column at a second mid-column feed
position. A distillation vapor stream is withdrawn from the
fractionation column below the mid-column feed positions and
directed in heat exchange relation with the LNG feed stream,
cooling the distillation vapor stream as it supplies at least part
of the heating of the LNG feed stream. The distillation vapor
stream is cooled sufficiently to condense at least a part of it,
forming a first condensed stream. At least a portion of the first
condensed stream is directed to the fractionation column at an
upper mid-column feed position. A portion of the column overhead
stream is also directed in heat exchange relation with the LNG feed
stream, so that it also supplies at least part of the heating of
the LNG feed stream as it is condensed, forming a second condensed
stream. The second condensed 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 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: |
43067387 |
Appl. No.: |
12/466661 |
Filed: |
May 15, 2009 |
Current U.S.
Class: |
62/620 |
Current CPC
Class: |
F25J 2210/02 20130101;
F25J 2210/62 20130101; F25J 3/0238 20130101; F25J 2200/02 20130101;
F25J 2200/30 20130101; F25J 2200/38 20130101; F25J 2290/40
20130101; F25J 3/0214 20130101; F25J 2230/08 20130101; F25J 3/0233
20130101; F25J 2210/06 20130101; F25J 2270/904 20130101; F25J
2200/72 20130101; F25J 2230/60 20130101; F25J 2205/04 20130101;
F25J 2240/02 20130101; F25J 2290/50 20130101; F25J 2235/60
20130101; F25J 3/0209 20130101; F25J 2200/76 20130101; F25J 3/0615
20130101 |
Class at
Publication: |
62/620 |
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 heated sufficiently to vaporize it,
thereby forming a vapor stream; (b) said vapor stream is expanded
to lower pressure and is thereafter supplied to a distillation
column at a first mid-column feed position; (c) said gas stream is
expanded to said lower pressure, is cooled, and is thereafter
supplied to said distillation column at a second mid-column feed
position; (d) a distillation vapor stream is withdrawn from a
region of said distillation column below said expanded vapor stream
and said expanded cooled gas stream, whereupon said distillation
vapor stream is cooled sufficiently to at least partially condense
it and form thereby a first condensed stream, with said cooling
supplying at least a portion of said heating of said liquefied
natural gas; (e) at least a portion of said first condensed stream
is supplied to said distillation column at an upper mid-column feed
position; (f) 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; (g) said compressed first portion
is cooled sufficiently to at least partially condense it and form
thereby a second condensed stream, with said cooling supplying at
least a portion of said heating of said liquefied natural gas; (h)
said second condensed stream is divided into at least a volatile
liquid stream and a reflux stream; (i) said reflux stream is
further cooled, with said cooling supplying at least a portion of
said heating of said liquefied natural gas; (j) said further cooled
reflux stream is supplied to said distillation column at a top
column feed position; (k) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of said expanded gas stream; (l) said
second portion is heated, with said heating supplying at least a
portion of said cooling of said expanded gas stream; (m) 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 (n) 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 heated sufficiently to partially
vaporize it; (b) said partially vaporized liquefied natural gas is
separated thereby to provide a vapor stream and a liquid stream;
(c) said vapor stream is expanded to lower pressure and is
thereafter supplied to a distillation column at a first mid-column
feed position; (d) said liquid stream is expanded to said lower
pressure and thereafter supplied to said distillation column at a
lower mid-column feed position; (e) said gas stream is expanded to
said lower pressure, is cooled, and is thereafter supplied to said
distillation column at a second mid-column feed position; (f) a
distillation vapor stream is withdrawn from a region of said
distillation column below said expanded vapor stream and said
expanded cooled gas stream, whereupon said distillation vapor
stream is cooled sufficiently to at least partially condense it and
form thereby a first condensed stream, with said cooling supplying
at least a portion of said heating of said liquefied natural gas;
(g) at least a portion of said first condensed stream is supplied
to said distillation column at an upper mid-column feed position;
(h) 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; (i) said compressed first portion is
cooled sufficiently to at least partially condense it and form
thereby a second condensed stream, with said cooling supplying at
least a portion of said heating of said liquefied natural gas; (j)
said second condensed stream is divided into at least a volatile
liquid stream and a reflux stream; (k) said reflux stream is
further cooled, with said cooling supplying at least a portion of
said heating of said liquefied natural gas; (l) said further cooled
reflux stream is supplied to said distillation column at a top
column feed position; (m) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of said expanded gas stream; (n) said
second portion is heated, with said heating supplying at least a
portion of said cooling of said expanded gas stream; (o) 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 (p) 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 heated sufficiently to vaporize it,
thereby forming a first vapor stream; (b) said first vapor stream
is expanded to lower pressure and is thereafter supplied to a
distillation column at a first mid-column feed position; (c) said
gas stream is expanded to said lower pressure and is thereafter
cooled sufficiently to partially condense it; (d) said partially
condensed gas stream is separated thereby to provide a second vapor
stream and a liquid stream; (e) said second vapor stream is further
cooled and thereafter supplied to said distillation column at a
second mid-column feed position; (f) said liquid stream is supplied
to said distillation column at a lower mid-column feed position;
(g) a distillation vapor stream is withdrawn from a region of said
distillation column below said expanded first vapor stream and said
further cooled second vapor stream, whereupon said distillation
vapor stream is cooled sufficiently to at least partially condense
it and form thereby a first condensed stream, with said cooling
supplying at least a portion of said heating of said liquefied
natural gas; (h) at least a portion of said first condensed stream
is supplied to said distillation column at an upper 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 second condensed stream, with said cooling supplying at
least a portion of said heating of said liquefied natural gas; (k)
said second 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 liquefied natural gas; (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 said expanded gas stream; (o) said
second portion is heated, with said heating supplying at least a
portion of said cooling of said expanded gas 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.
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 heated sufficiently to partially
vaporize it; (b) said partially vaporized liquefied natural gas is
separated thereby to provide a first vapor stream and a first
liquid stream; (c) said first vapor stream is expanded to lower
pressure and is thereafter supplied to a distillation column at a
first mid-column feed position; (d) said first liquid 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 expanded to said lower pressure and is
thereafter cooled sufficiently to partially condense it; (f) said
partially condensed gas stream is separated thereby to provide a
second vapor stream and a second liquid stream; (g) said second
vapor stream is further cooled and thereafter supplied to said
distillation column at a second mid-column feed position; (h) said
second liquid stream is supplied to said distillation column at a
second lower mid-column feed position; (i) a distillation vapor
stream is withdrawn from a region of said distillation column below
said expanded first vapor stream and said further cooled second
vapor stream, whereupon said distillation vapor stream is cooled
sufficiently to at least partially condense it and form thereby a
first condensed stream, with said cooling supplying at least a
portion of said heating of said liquefied natural gas; (j) at least
a portion of said first condensed stream is supplied to said
distillation column at an upper 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
second condensed stream, with said cooling supplying at least a
portion of said heating of said liquefied natural gas; (m) said
second 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 liquefied natural gas; (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 said expanded gas stream; (q) said
second portion is heated, with said heating supplying at least a
portion of said cooling of said expanded gas 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.
5. The process according to claim 1 or 2 wherein (a) said gas
stream is cooled, is expanded to said lower pressure, and is
thereafter supplied to said distillation column at said second
mid-column feed position; (b) said distillation vapor stream is
withdrawn from a region of said distillation column below said
expanded vapor stream and said cooled expanded gas stream; (c) said
volatile liquid stream is heated sufficiently to vaporize it, with
said heating supplying at least a portion of said cooling of said
gas stream; and (d) said second portion is heated, with said
heating supplying at least a portion of said cooling of said gas
stream.
6. The process according to claim 3 wherein (a) said gas stream is
cooled sufficiently to partially condense it; thereby forming said
second vapor stream and said liquid stream; (b) said second vapor
stream is expanded to said lower pressure and is thereafter
supplied to said distillation column at said second mid-column feed
position; (c) said liquid stream is expanded to said lower pressure
and is thereafter supplied to said distillation column at said
lower mid-column feed position; (d) said distillation vapor stream
is withdrawn from a region of said distillation column below said
expanded first vapor stream and said expanded second vapor stream;
(e) said volatile liquid stream is heated sufficiently to vaporize
it, with said heating supplying at least a portion of said cooling
of said gas stream; and (f) said second portion is heated, with
said heating supplying at least a portion of said cooling of said
gas stream.
7. The process according to claim 4 wherein (a) said gas stream is
cooled sufficiently to partially condense it; thereby forming said
second vapor stream and said second liquid stream; (b) said second
vapor stream is expanded to said lower pressure and is thereafter
supplied to said distillation column at said second mid-column feed
position; (c) said second liquid stream is expanded to said lower
pressure and is thereafter supplied to said distillation column at
said second lower mid-column feed position; (d) said distillation
vapor stream is withdrawn from a region of said distillation column
below said expanded first vapor stream and said expanded second
vapor stream; (e) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of said gas stream; and (f) said second
portion is heated, with said heating supplying at least a portion
of said cooling of said gas stream.
8. The process according to claim 1, 2, 3, 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 said expanded gas stream; and
(c) said vaporized volatile liquid stream and said heated
compressed second portion are combined to form said volatile
residue gas fraction.
9. 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 heated sufficiently to vaporize it,
thereby forming a vapor stream; (b) said vapor stream is expanded
to lower pressure and is thereafter supplied at a first lower feed
position to an absorber column that produces an overhead
distillation stream and a bottom liquid stream; (c) said gas stream
is expanded to said lower pressure, is cooled, and is thereafter
supplied to said absorber column at a second lower feed position;
(d) said bottom liquid stream is supplied at a top column feed
position to a stripper column that produces an overhead vapor
stream and said relatively less volatile liquid fraction; (e) said
overhead vapor stream is divided into at least a first distillation
vapor stream and a second distillation vapor stream, whereupon said
second distillation vapor stream is supplied to said absorber
column at a third lower feed position; (f) said first distillation
vapor stream is cooled sufficiently to at least partially condense
it and form thereby a first condensed stream, with said cooling
supplying at least a portion of said heating of said liquefied
natural gas; (g) at least a portion of said first condensed stream
is supplied to said absorber column at a mid-column feed position;
(h) said overhead distillation stream is divided into at least a
first portion and a second portion, whereupon said first portion is
compressed to higher pressure; (i) said compressed first portion is
cooled sufficiently to at least partially condense it and form
thereby a second condensed stream, with said cooling supplying at
least a portion of said heating of said liquefied natural gas; (j)
said second condensed stream is divided into at least a volatile
liquid stream and a reflux stream; (k) said reflux stream is
further cooled, with said cooling supplying at least a portion of
said heating of said liquefied natural gas; (l) said further cooled
reflux stream is supplied to said absorber column at a top column
feed position; (m) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of said expanded gas stream; (n) said
second portion is heated, with said heating supplying at least a
portion of said cooling of said expanded gas stream; (o) 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 (p) the quantity and temperature
of said reflux stream and the temperatures of said feeds to said
absorber column and said stripper column are effective to maintain
the overhead temperatures of said absorber column and said stripper
column at temperatures whereby the major portion of said heavier
hydrocarbon components is recovered in said relatively less
volatile liquid fraction by fractionation in said absorber column
and said stripper column.
10. 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 heated sufficiently to partially
vaporize it; (b) said partially vaporized liquefied natural gas is
separated thereby to provide a vapor stream and a liquid stream;
(c) said vapor stream is expanded to lower pressure and is
thereafter supplied at a first lower feed position to an absorber
column that produces an overhead distillation stream and a bottom
liquid stream; (d) said gas stream is expanded to said lower
pressure, is cooled, and is thereafter supplied to said absorber
column at a second lower feed position; (e) said bottom liquid
stream is supplied at a top column feed position to a stripper
column that produces an overhead vapor stream and said relatively
less volatile liquid fraction; (f) said liquid stream is expanded
to said lower pressure and thereafter supplied to said stripper
column at a mid-column feed position; (g) said overhead vapor
stream is divided into at least a first distillation vapor stream
and a second distillation vapor stream, whereupon said second
distillation vapor stream is supplied to said absorber column at a
third lower feed position; (h) said first distillation vapor stream
is cooled sufficiently to at least partially condense it and form
thereby a first condensed stream, with said cooling supplying at
least a portion of said heating of said liquefied natural gas; (i)
at least a portion of said first condensed stream is supplied to
said absorber column at a mid-column feed position; (j) said
overhead distillation stream is divided into at least a first
portion and a second portion, whereupon said first portion is
compressed to higher pressure; (k) said compressed first portion is
cooled sufficiently to at least partially condense it and form
thereby a second condensed stream, with said cooling supplying at
least a portion of said heating of said liquefied natural gas; (l)
said second condensed stream is divided into at least a volatile
liquid stream and a reflux stream; (m) said reflux stream is
further cooled, with said cooling supplying at least a portion of
said heating of said liquefied natural gas; (n) said further cooled
reflux stream is supplied to said absorber column at a top column
feed position; (o) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of said expanded gas stream; (p) said
second portion is heated, with said heating supplying at least a
portion of said cooling of said expanded gas stream; (q) 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 (r) the quantity and temperature
of said reflux stream and the temperatures of said feeds to said
absorber column and said stripper column are effective to maintain
the overhead temperatures of said absorber column and said stripper
column at temperatures whereby the major portion of said heavier
hydrocarbon components is recovered in said relatively less
volatile liquid fraction by fractionation in said absorber column
and said stripper column.
11. 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 heated sufficiently to vaporize it,
thereby forming a first vapor stream; (b) said first vapor stream
is expanded to lower pressure and is thereafter supplied at a first
lower feed position to an absorber column that produces an overhead
distillation stream and a bottom liquid stream; (c) said gas stream
is expanded to said lower pressure and is thereafter cooled
sufficiently to partially condense it; (d) said partially condensed
gas stream is separated thereby to provide a second vapor stream
and a liquid stream; (e) said second vapor stream is further cooled
and thereafter supplied to said absorber column at a second lower
feed position; (f) said bottom liquid stream is supplied at a top
column feed position to a stripper column that produces an overhead
vapor stream and said relatively less volatile liquid fraction; (g)
said liquid stream is supplied to said stripper column at a
mid-column feed position; (h) said overhead vapor stream is divided
into at least a first distillation vapor stream and a second
distillation vapor stream, whereupon said second distillation vapor
stream is supplied to said absorber column at a third lower feed
position; (i) said first distillation vapor stream is cooled
sufficiently to at least partially condense it and form thereby a
first condensed stream, with said cooling supplying at least a
portion of said heating of said liquefied natural gas; (j) at least
a portion of said first condensed stream is supplied to said
absorber column at a mid-column feed position; (k) said overhead
distillation stream is 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
second condensed stream, with said cooling supplying at least a
portion of said heating of said liquefied natural gas; (m) said
second 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 liquefied natural gas; (o) said further cooled
reflux stream is supplied to said absorber 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 said expanded gas stream; (q) said
second portion is heated, with said heating supplying at least a
portion of said cooling of said expanded gas 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
absorber column and said stripper column are effective to maintain
the overhead temperatures of said absorber column and said stripper
column at temperatures whereby the major portion of said heavier
hydrocarbon components is recovered in said relatively less
volatile liquid fraction by fractionation in said absorber column
and said stripper column.
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 heated sufficiently to partially
vaporize it; (b) said partially vaporized liquefied natural gas is
separated thereby to provide a first vapor stream and a first
liquid stream; (c) said first vapor stream is expanded to lower
pressure and is thereafter supplied at a first lower feed position
to an absorber column that produces an overhead distillation stream
and a bottom liquid stream; (d) said gas stream is expanded to said
lower pressure and is thereafter cooled sufficiently to partially
condense it; (e) said partially condensed gas stream is separated
thereby to provide a second vapor stream and a second liquid
stream; (f) said second vapor stream is further cooled and
thereafter supplied to said absorber column at a second lower feed
position; (g) said bottom liquid stream is supplied at a top column
feed position to a stripper column that produces an overhead vapor
stream and said relatively less volatile liquid fraction; (h) said
first liquid stream is expanded to said lower pressure and
thereafter supplied to said stripper column at a first mid-column
feed position; (i) said second liquid stream is supplied to said
stripper column at a second mid-column feed position; (j) said
overhead vapor stream is divided into at least a first distillation
vapor stream and a second distillation vapor stream, whereupon said
second distillation vapor stream is supplied to said absorber
column at a third lower feed position; (k) said first distillation
vapor stream is cooled sufficiently to at least partially condense
it and form thereby a first condensed stream, with said cooling
supplying at least a portion of said heating of said liquefied
natural gas; (l) at least a portion of said first condensed stream
is supplied to said absorber column at a mid-column feed position;
(m) said overhead distillation stream is divided into at least a
first portion and a second portion, whereupon said first portion is
compressed to higher pressure; (n) said compressed first portion is
cooled sufficiently to at least partially condense it and form
thereby a second condensed stream, with said cooling supplying at
least a portion of said heating of said liquefied natural gas; (o)
said second condensed stream is divided into at least a volatile
liquid stream and a reflux stream; (p) said reflux stream is
further cooled, with said cooling supplying at least a portion of
said heating of said liquefied natural gas; (q) said further cooled
reflux stream is supplied to said absorber column at a top column
feed position; (r) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of said expanded gas stream; (s) said
second portion is heated, with said heating supplying at least a
portion of said cooling of said expanded gas stream; (t) 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 (u) the quantity and temperature
of said reflux stream and the temperatures of said feeds to said
absorber column and said stripper column are effective to maintain
the overhead temperatures of said absorber column and said stripper
column at temperatures whereby the major portion of said heavier
hydrocarbon components is recovered in said relatively less
volatile liquid fraction by fractionation in said absorber column
and said stripper column.
13. The process according to claim 9 or 10 wherein (a) said gas
stream is cooled, is expanded to said lower pressure, and is
thereafter supplied to said absorber column at said second lower
feed position; (b) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of said gas stream; and (c) said second
portion is heated, with said heating supplying at least a portion
of said cooling of said gas stream.
14. The process according to claim 11 wherein (a) said gas stream
is cooled sufficiently to partially condense it; thereby forming
said second vapor stream and said liquid stream; (b) said second
vapor stream is expanded to said lower pressure and is thereafter
supplied to said absorber column at said second lower feed
position; (c) said liquid stream is expanded to said lower pressure
and is thereafter supplied to said stripper column at said
mid-column feed position; (d) said volatile liquid stream is heated
sufficiently to vaporize it, with said heating supplying at least a
portion of said cooling of said gas stream; and (e) said second
portion is heated, with said heating supplying at least a portion
of said cooling of said gas stream.
15. The process according to claim 12 wherein (a) said gas stream
is cooled sufficiently to partially condense it; thereby forming
said second vapor stream and said second liquid stream; (b) said
second vapor stream is expanded to said lower pressure and is
thereafter supplied to said absorber column at said second lower
feed position; (c) said second liquid stream is expanded to said
lower pressure and is thereafter supplied to said stripper column
at said second mid-column feed position; (d) said volatile liquid
stream is heated sufficiently to vaporize it, with said heating
supplying at least a portion of said cooling of said gas stream;
and (e) said second portion is heated, with said heating supplying
at least a portion of said cooling of said gas stream.
16. The process according to claim 9, 10, 11, or 12 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 said expanded gas stream; and
(c) said vaporized volatile liquid stream and said heated
compressed second portion are combined to form said volatile
residue gas fraction.
17. The process according to claim 1, 2, 3, 4, 6, 7, 9, 10, 11, 12,
14, or 15 wherein said volatile residue gas fraction contains a
major portion of said methane and C.sub.2 components.
18. The process according to claim 5 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
19. The process according to claim 8 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
20. The process according to claim 13 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
21. The process according to claim 16 wherein said volatile residue
gas fraction contains a major portion of said methane and C.sub.2
components.
Description
BACKGROUND OF THE INVENTION
[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.
[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 co-pending application Ser.
No. 11/430,412 could be 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 co-pending application Ser. No.
11/430,412, 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. No. 7,216,507 and
co-pending application Ser. No. 11/430,412, 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 23 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 24 wherein the
condensed liquid (stream 46) is separated from any uncondensed
vapor (stream 45). The liquid stream 46 from reflux separator 24 is
pumped by reflux pump 25 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 24 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 [9,404 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 23 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. No. 7,216,507 and co-pending application Ser. No.
11/430,412, 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 [9,404 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 [9,198 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
1278 psia [8,809 kPa(a)] for subsequent vaporization in heat
exchangers 14 and 12, heating stream 83a to -114.degree. F.
[-81.degree. C.] and then to 40.degree. F. [4.degree. C.] as
described in paragraphs [0035] and [0032] below to produce warm
lean LNG stream 83c.
[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 cool lean LNG (stream 83b), cool
overhead distillation stream 38a at -114.degree. F. [-81.degree.
C.], and demethanizer liquids (stream 39) at -51.degree. F.
[-46.degree. C.]. The cooled stream 31a enters separator 13 at
-91.degree. F. [-68.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 390 psia [2,687 kPa(a)]) of
fractionation tower 20. The expanded stream 35a leaving expansion
valve 17 reaches a temperature of -111.degree. F. [-80.degree. C.]
and is supplied to fractionation tower 20 at a first lower
mid-column feed point.
[0033] Vapor stream 34 from separator 13 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 -121.degree. F. [-85.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.
[0034] 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 89.degree.
F. [31.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).
[0035] A portion of the distillation vapor (stream 44) is withdrawn
from the upper region of the stripping section of fractionation
column 20 at -125.degree. F. [-87.degree. C.] and compressed to 545
psia [3,756 kPa(a)] by compressor 26. The compressed stream 44a is
then cooled from -87.degree. F. [-66.degree. C.] to -143.degree. F.
[-97.degree. C.] and condensed (stream 44b) in heat exchanger 14 by
heat exchange with cold overhead distillation stream 38 exiting the
top of demethanizer 20 and cold lean LING (stream 83a) at
-116.degree. F. [-82.degree. C.]. Condensed liquid stream 44b is
expanded by expansion valve 16 to a pressure slightly above the
operating pressure of demethanizer 20, and the resulting stream 44c
at -146.degree. F. [-99.degree. C.] is then supplied as cold liquid
reflux to an intermediate region in the absorbing section of
demethanizer 20. This supplemental reflux absorbs and condenses
most of the C.sub.3 components and heavier components (as well as
some of the C.sub.2 components) from the vapors rising in the lower
rectification region of the absorbing section so that only a small
amount of recycle (stream 36) must be cooled, condensed, subcooled,
and flash expanded to produce the top reflux stream 36c that
provides the final rectification in the upper region of the
absorbing section of demethanizer 20. As the cold reflux stream 36c
contacts the rising vapors in the upper region of the absorbing
section, it condenses and absorbs the C.sub.2 components and any
remaining C.sub.3 components and heavier components from the vapors
so that they can be captured in the bottom product (stream 40) from
demethanizer 20.
[0036] Overhead distillation stream 38 is withdrawn from the upper
section of fractionation tower 20 at -148.degree. F. [-100.degree.
C.]. It passes countercurrently to compressed distillation vapor
stream 44a and recycle stream 36a in heat exchanger 14 where it is
heated to -114.degree. F. [-81 .degree. C.] (stream 38a), and
countercurrently to inlet gas stream 31 and recycle stream 36 in
heat exchanger 12 where it is heated to 20.degree. F. [-7.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 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.
[0037] Recycle stream 36 flows to heat exchanger 12 and is cooled
to -105.degree. F. [-76.degree. C.] by heat exchange with cool lean
LNG (stream 83b), cool overhead distillation stream 38a, and
demethanizer liquids (stream 39) as described previously. Stream
36a is further cooled to -143.degree. F. [-97.degree. C.] by heat
exchange with cold lean LNG stream 83a and cold overhead
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 -151.degree. F. [-102.degree. C.]. The expanded
stream 36c is then supplied to fractionation tower 20 as the top
column feed. Any vapor portion of stream 36c combines with the
vapors rising from the top fractionation stage of the column to
form overhead distillation stream 38, which is withdrawn from an
upper region of the tower as described previously.
[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 34 28,762 1,051 163 22 30,759 35
13,783 3,997 2,809 1,636 22,386 44 6,746 195 3 0 7,000 38 49,040 39
0 0 50,064 36 6,595 5 0 0 6,733 37 42,445 34 0 0 43,331 40 100
5,014 2,972 1,658 9,814 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 48 0 0 83,844 41 153 7,642 3,463 1,661 12,990 Recoveries*
Ethane 99.38% Propane 100.00% Butanes+ 100.00% Power LNG Feed Pump
3,552 HP [5,839 kW] LNG Product Pump 1,774 HP [2,916 kW] Residue
Gas Compressor 29,272 HP [48,123 kW] Reflux Compressor 601 HP [988
kW] Totals 35,199 HP [57,866 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 26,780 MBTU/Hr [17,298 kW]
Demethanizer Reboiler 61 3,400 MBTU/Hr [2,196 kW] Totals 30,180
MBTU/Hr [19,494 kW] Specific Power HP-Hr/Lb. Mole 2.710 [kW-Hr/kg
mole] [4.455] *(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
99.38%, 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 more
than 5% 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 [9.404 kPa(a)] and
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 71a is
first heated to -54.degree. F. [-48.degree. C.] in heat exchanger
52 by cooling compressed distillation stream 81a at -32.degree. F.
[-36.degree. C.], reflux stream 82, and distillation vapor stream
44. The partially heated stream 71b 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 19, 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.)
[0042] The heated stream 71c enters separator 54 at 11.degree. F.
[-12.degree. C.] and 1334 psia [9,198 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 the high pressure feed. The
machine 55 expands the vapor substantially isentropically to the
tower operating pressure (approximately 412 psia [2,839 kPa(a)]),
with the work expansion cooling the expanded stream 77a to a
temperature of approximately -100.degree. F. [-73.degree. C.]. The
work recovered is often used to drive a centrifugal compressor
(such as item 56) that can be used to re-compress a portion (stream
81) of the column overhead vapor (stream 79), for example. The
partially condensed expanded stream 77a is thereafter supplied as
feed to fractionation column 20 at a first 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 first lower mid-column feed point.
[0043] 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 cooled in heat
exchanger 12 by heat exchange with cool lean LNG (stream 83a) at
-99.degree. F. [-73.degree. C.], cold distillation stream 38, and
demethanizer liquids (stream 39) at -57.degree. F. [-50.degree.
C.]. The cooled stream 31a enters separator 13 at -82.degree. F.
[-63.degree. C.] and 584 psia [4,027 kPa(a)] where the vapor
(stream 34) is separated from the condensed liquid (stream 35).
Note that in all cases exchanger 12 is 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.)
[0044] The vapor (stream 34) from separator 13 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 operating pressure of
fractionation tower 20, with the work expansion cooling the
expanded stream 34a to a temperature of approximately -108.degree.
F. [-78.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 38a), for
example. The expanded partially condensed stream 34a is supplied to
fractionation tower 20 at a second 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 -99.degree. F. [-73.degree. C.]
and is supplied to fractionation tower 20 at a second lower
mid-column feed point.
[0045] 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,
side reboiler 18 using low level utility heat, 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 41 exits the
bottom of the tower at 83.degree. F. [28.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.
[0046] A portion of the distillation vapor (stream 44) is withdrawn
from the upper region of stripping section 20b of fractionation
column 20 at -120.degree. F. [-84.degree. C.] and is cooled to
-143.degree. F. [-97.degree. C.] and condensed (stream 44a) in heat
exchanger 52 by heat exchange with the cold LNG (stream 71a).
Condensed liquid stream 44a is pumped to slightly above the
operating pressure of fractionation column 20 by pump 27, whereupon
stream 44b at -143.degree. F. [-97.degree. C.] is then supplied as
cold liquid reflux to an intermediate region in absorbing section
20a of fractionation column 20. This supplemental reflux absorbs
and condenses most of the C.sub.3 components and heavier components
(as well as some of the C.sub.2 components) from the vapors rising
in the lower rectification region of absorbing section 20a so that
only a small amount of the lean LNG (stream 82) must be subcooled
to produce the top reflux stream 82b that provides the final
rectification in the upper region of absorbing section 20a of
fractionation column 20.
[0047] Overhead distillation stream 79 is withdrawn from the upper
section of fractionation tower 20 at -145.degree. F. [-98.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 1092 psia [7,529
kPa(a)] (stream 81a). At this pressure, the stream is totally
condensed as it is cooled to -106.degree. F. [-77.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 1273 psia [8,777 kPa(a)] for
subsequent vaporization in heat exchanger 12, heating stream 83a to
65.degree. F. [18.degree. C.] as described previously to produce
warm lean LNG stream 83b.
[0048] The remaining portion of stream 81b (stream 82) flows to
heat exchanger 52 where it is subcooled to -234.degree. F.
[-148.degree. C.] by heat exchange with the cold LNG (stream 71a)
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 -232.degree. F. [-146.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 region of absorbing
section 20a of demethanizer 20.
[0049] The second portion of overhead distillation stream 79
(stream 38) flows countercurrently to inlet gas stream 31 in heat
exchanger 12 where it is heated to -62.degree. F. [-52.degree. C.]
(stream 38a). 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 38b to sales gas
line pressure (stream 38c). (Note that discharge cooler 22 is not
needed in this example. Some applications may require cooling of
compressed distillation stream 38c so that the resultant
temperature when mixed with warm lean LNG stream 83b is
sufficiently cool to comply with the requirements of the sales gas
pipeline.) Stream 38c/38d then combines with warm lean LNG stream
83b to form the residue gas product (stream 42). Residue gas stream
42 at 89.degree. F. [32.degree. C.] flows to the sales gas pipeline
at 1262 psia [8,701 kPa(a)], sufficient to meet line
requirements.
[0050] 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 34 32,557 1,468 247 35 35,112 35
9,988 3,580 2,725 1,623 18,033 71 40,293 2,642 491 3 43,689 77
40,293 2,642 491 3 43,689 78 0 0 0 0 0 44 23,473 771 21 0 24,399 79
91,871 58 0 0 93,147 38 55,581 35 0 0 56,354 81 36,290 23 0 0
36,793 82 9,186 6 0 0 9,313 83 27,104 17 0 0 27,480 42 82,685 52 0
0 83,834 41 153 7,638 3,463 1,661 13,000 Recoveries* Ethane 99.33%
Propane 100.00% Butanes+ 100.00% Power LNG Feed Pump 3,552 HP
[5,839 kW] LNG Product Pump 569 HP [935 kW] Reflux Pump 87 HP [143
kW] Residue Gas Compressor 22,960 HP [37,746 kW] Totals 27,168 HP
[44,663 kW] Low Level Utility Heat Liquid Feed Heater 58,100
MBTU/Hr [37,530 kW] Demethanizer Reboiler 18 28,000 MBTU/Hr [5,167
kW] Totals 66,100 MBTU/Hr [42,697 kW] High Level Utility Heat
Demethanizer Reboiler 19 31,130 MBTU/Hr [20,108 kW] Specific Power
HP-Hr/Lb. Mole 2.090 [kW-Hr/kg mole] [3.436] *(Based on un-rounded
flow rates)
[0051] 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 99.33%, 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 the power required for
the FIG. 3 embodiment of the present invention is nearly 4% lower
than the FIG. 1 process, meaning that 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
2.090 HP-Hr/Lb. Mole [3.436 kW-Hr/kg mole] for the FIG. 3
embodiment of the present invention, an increase of more than 27%
in the production efficiency. In addition, the high level utility
heat requirement for the FIG. 3 embodiment of the present invention
is only 39% of the requirement for the FIG. 1 process.
[0052] 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 nearly 23% lower than the
FIG. 2 processes. This results in reducing the specific power from
2.710 HP-Hr/Lb. Mole [4.455 kW-Hr/kg mole] for the FIG. 2 processes
to 2.090 HP-Hr/Lb. Mole [3.436 kW-Hr/kg mole] for the FIG. 3
embodiment of the present invention, an improvement of nearly 23%
in the production efficiency.
[0053] There are five 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 20. 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 upper region of absorbing section
20a in fractionation tower 20 and avoiding the equilibrium
limitations of such prior art processes. Second, using distillation
vapor stream 44 to produce supplemental reflux for the lower region
of absorbing section 20a in fractionation column 20 allows using
less top reflux (stream 82b) for fractionation tower 20. 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 20, reducing the duty required in
reboiler 19 and minimizing the amount of high level utility heat
needed to meet the specification for the bottom liquid product from
demethanizer 20. Third, the rectification of the column vapors
provided by absorbing section 20a allows all of the LNG feed to be
vaporized before entering work expansion machine 55 as stream 77,
resulting in significant power recovery. This power can then be
used to compress the first portion (stream 81) of distillation
overhead stream 79 to a pressure sufficiently high so that it can
be condensed in heat exchanger 52 and so that the resulting lean
LNG (stream 83) can then be pumped to the pipeline delivery
pressure. (Pumping uses significantly less power than
compressing.)
[0054] Fourth, using the cold lean LNG stream 83a to provide "free"
refrigeration to the gas stream in heat exchanger 12 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, this "free" refrigeration of inlet
gas stream 31 means less of the cooling duty in heat exchanger 12
must be supplied by distillation vapor stream 38, so that stream
38a is cooler and less compression power is needed to raise its
pressure to the pipeline delivery condition.
Example 2
[0055] An alternative method of processing LNG and 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.
[0056] 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 [9,404 kPa(a)] and
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 71a is
first heated to -66.degree. F. [-54.degree. C.] in heat exchanger
52 by cooling compressed distillation stream 81a at -54.degree. F.
[-48.degree. C.], reflux stream 82, and distillation vapor stream
44. The partially heated stream 71b is further heated in heat
exchanger 53 using low level utility heat.
[0057] The heated stream 71c enters separator 54 at 3.degree. F.
[-16.degree. C.] and 1334 psia [9,198 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 the high pressure feed. The
machine 55 expands the vapor substantially isentropically to the
tower operating pressure (approximately 420 psia [2,896 kPa(a)]),
with the work expansion cooling the expanded stream 77a to a
temperature of approximately -102.degree. F. [-75.degree. C.]. The
partially condensed expanded stream 77a is thereafter supplied as
feed to fractionation column 20 at a first 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 first lower mid-column feed point.
[0058] 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 enters a work expansion
machine 10 in which mechanical energy is extracted from 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 31a to a temperature of approximately 93.degree. F.
[34.degree. C.]. The expanded stream 31a is further cooled in heat
exchanger 12 by heat exchange with cool lean LNG (stream 83a) at
-93.degree. F. [-69.degree. C.], cool distillation stream 38a, and
demethanizer liquids (stream 39) at -76.degree. F. [-60.degree.
C.].
[0059] The cooled stream 31b enters separator 13 at -81.degree. F.
[-63.degree. C.] and 428 psia [2,949 kPa(a)] where the vapor
(stream 34) is separated from the condensed liquid (stream 35).
Vapor stream 34 is cooled to -122.degree. F. [-86.degree. C.] in
heat exchanger 14 by heat exchange with cold distillation stream
38, and the partially condensed stream 34a is then supplied to
fractionation tower 20 at a second mid-column feed point. Liquid
stream 35 is directed through valve 17 and is supplied to
fractionation tower 20 at a second lower mid-column feed point.
[0060] A portion of the distillation vapor (stream 44) is withdrawn
from the upper region of the stripping section of fractionation
column 20 at -119.degree. F. [-84.degree. C.] and is cooled to
-145.degree. F. [-98.degree. C.] and condensed (stream 44a) in heat
exchanger 52 by heat exchange with the cold LNG (stream 71a).
Condensed liquid stream 44a is pumped to slightly above the
operating pressure of fractionation column 20 by pump 27, whereupon
stream 44b at -144.degree. F. [-98.degree. C.] is then supplied as
cold liquid reflux to an intermediate region in the absorbing
section of fractionation column 20. This supplemental reflux
absorbs and condenses most of the C.sub.3 components and heavier
components (as well as some of the C.sub.2 components) from the
vapors rising in the lower rectification region of the absorbing
section of fractionation column 20.
[0061] The column liquid stream 41 exits the bottom of the tower at
85.degree. F. [29.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 -144.degree. F.
[-98.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 929 psia
[6,405 kPa(a)] (stream 81a). At this pressure, the stream is
totally condensed as it is cooled to -108.degree. F. [-78.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 1273 psia [8,777 kPa(a)] for
subsequent vaporization in heat exchanger 12, heating stream 83a to
65.degree. F. [18.degree. C.] as described previously to produce
warm lean LNG stream 83b.
[0062] The remaining portion of stream 81b (stream 82) flows to
heat exchanger 52 where it is subcooled to -235.degree. F.
[-148.degree. C.] by heat exchange with the cold LNG (stream 71a)
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 -233.degree. F. [-147.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 region of the absorbing
section of demethanizer 20.
[0063] The second portion of overhead distillation stream 79
(stream 38) flows countercurrently to separator vapor stream 34 in
heat exchanger 14 where it is heated to -87.degree. F. [-66.degree.
C.] (stream 38a), and to expanded inlet gas stream 31a in heat
exchanger 12 where it is heated to -47.degree. F. [-44.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). Stream 38d/38e then combines with warm
lean LNG stream 83b to form the residue gas product (stream 42).
Residue gas stream 42 at 99.degree. F. [37.degree. C.] flows to the
sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient to meet
line requirements.
[0064] 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 34 37,612 2,081 327 39 40,922 35
4,933 2,967 2,645 1,619 12,223 71 40,293 2,642 491 3 43,689 77
40,293 2,642 491 3 43,689 78 0 0 0 0 0 44 15,646 515 14 0 16,250 79
92,556 62 0 0 93,856 38 48,684 32 0 0 49,369 81 43,872 30 0 0
44,487 82 9,871 7 0 0 10,010 83 34,001 23 0 0 34,477 42 82,685 55 0
0 83,846 41 153 7,635 3,463 1,661 12,988 Recoveries* Ethane 99.29%
Propane 100.00% Butanes+ 100.00% Power LNG Feed Pump 3,552 HP
[5,839 kW] LNG Product Pump 1,437 HP [2,363 kW] Reflux Pump 58 HP
[95 kW] Residue Gas Compressor 18,325 HP [30,126 kW] Totals 23,372
HP [38,423 kW] Low Level Utility Heat Liquid Feed Heater 66,000
MBTU/Hr [42,632 kW] Demethanizer Reboiler 18 17,300 MBTU/Hr [11,175
kW] Totals 83,300 MBTU/Hr [53,807 kW] High Level Utility Heat
Demethanizer Reboiler 19 32,940 MBTU/Hr [21,278 kW] Specific Power
HP-Hr/Lb. Mole 1.800 [kW-Hr/kg mole] [2.958] *(Based on un-rounded
flow rates)
[0065] 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 nearly 14%. However, the high level utility
heat required for the FIG. 4 embodiment of the present invention is
slightly higher (about 6%) than that of the FIG. 3 embodiment.
Example 3
[0066] Another alternative method of processing LNG and 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.
[0067] 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 [9,404 kPa(a)] and
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 71a is
first heated to -71.degree. F. [-57.degree. C.] in heat exchanger
52 by cooling compressed distillation stream 81a at -25.degree. F.
[-32.degree. C.], reflux stream 82, distillation vapor stream 44,
and separator vapor stream 34. The partially heated stream 71b is
further heated in heat exchanger 53 using low level utility
heat.
[0068] The heated stream 71c enters separator 54 at 1.degree. F.
[-17.degree. C.] and 1334 psia [9,198 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 the high pressure feed. The
machine 55 expands the vapor substantially isentropically to the
tower operating pressure (approximately 395 psia [2,721 kPa(a)]),
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 20 at a first 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 first lower mid-column feed point.
[0069] 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 enters a work expansion
machine 10 in which mechanical energy is extracted from 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 31a to a temperature of approximately 87.degree. F.
[30.degree. C.]. The expanded stream 31a is further cooled in heat
exchanger 12 by heat exchange with cool lean LNG (stream 83a) at
-97.degree. F. [-72.degree. C.], cool distillation stream 38b, and
demethanizer liquids (stream 39) at -81.degree. F. [-63.degree.
C.].
[0070] The cooled stream 31b enters separator 13 at -81.degree. F.
[-63.degree. C.] and 403 psia [2,777 kPa(a)] where the vapor
(stream 34) is separated from the condensed liquid (stream 35).
Vapor stream 34 is cooled to -117.degree. F. [-83.degree. C.] in
heat exchanger 52 by heat exchange with cold LNG stream 71a and
compressed distillation stream 38a, and the partially condensed
stream 34a is then supplied to fractionation tower 20 at a second
mid-column feed point. Liquid stream 35 is directed through valve
17 and is supplied to fractionation tower 20 at a second lower
mid-column feed point.
[0071] A portion of the distillation vapor (stream 44) is withdrawn
from the upper region of the stripping section of fractionation
column 20 at -119.degree. F. [-84.degree. C.] and is cooled to
-145.degree. F. [-98.degree. C.] and condensed (stream 44a) in heat
exchanger 52 by heat exchange with the cold LNG (stream 71a).
Condensed liquid stream 44a is pumped to slightly above the
operating pressure of fractionation column 20 by pump 27, whereupon
stream 44b at -144.degree. F. [-98.degree. C.] is then supplied as
cold liquid reflux to an intermediate region in the absorbing
section of fractionation column 20. This supplemental reflux
absorbs and condenses most of the C.sub.3 components and heavier
components (as well as some of the C.sub.2 components) from the
vapors rising in the lower rectification region of the absorbing
section of fractionation column 20.
[0072] The column liquid stream 41 exits the bottom of the tower at
79.degree. F. [26.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 -147.degree. F.
[-99.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 1124 psia
[7,750 kPa(a)] (stream 81a). At this pressure, the stream is
totally condensed as it is cooled to -103.degree. F. [-75.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 1273 psia [8,777 kPa(a)] for
subsequent vaporization in heat exchanger 12, heating stream 83a to
65.degree. F. [18.degree. C.] as described previously to produce
warm lean LNG stream 83b.
[0073] The remaining portion of stream 81b (stream 82) flows to
heat exchanger 52 where it is subcooled to -236.degree. F.
[-149.degree. C.] by heat exchange with the cold LNG (stream 71a)
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 -233.degree. F. [-147.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 region of the absorbing
section of demethanizer 20.
[0074] The second portion of overhead distillation stream 79
(stream 38) is compressed to 625 psia [4,309 kPa(a)] by compressor
11 driven by expansion machine 10. It then flows countercurrently
to separator vapor stream 34 in heat exchanger 52 where it is
heated from -97.degree. F. [-72.degree. C.] to -65.degree. F.
[-53.degree. C.] (stream 38b), an inlet gas stream 31a in heat
exchanger 12 where it is heated to 12.degree. F. [-11.degree. C.]
(stream 38c). The distillation stream is then further compressed to
sales gas line pressure (stream 38d) in compressor 21 driven by a
supplemental power source, and stream 38d/38e then combines with
warm lean LNG stream 83b to form the residue gas product (stream
42). Residue gas stream 42 at 107.degree. F. [42.degree. C.] flows
to the sales gas pipeline at 1262 psia [8,701 kPa(a)], sufficient
to meet line requirements.
[0075] 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 34 38,194 2,203 348 40 41,654 35
4,351 2,845 2,624 1,618 11,491 71 40,293 2,642 491 3 43,689 77
40,293 2,642 491 3 43,689 78 0 0 0 0 0 44 17,004 614 16 0 17,715 79
91,637 60 0 0 92,925 38 59,566 39 0 0 60,403 81 32,071 21 0 0
32,522 82 8,952 6 0 0 9,078 83 23,119 15 0 0 23,444 42 82,685 54 0
0 83,847 41 153 7,636 3,463 1,661 12,987 Recoveries* Ethane 99.30%
Propane 100.00% Butanes+ 100.00% Power LNG Feed Pump 3,552 HP
[5,839 kW] LNG Product Pump 418 HP [687 kW] Reflux Pump 63 HP [104
kW] Residue Gas Compressor 19,274 HP [31,686 kW] Totals 23,307 HP
[38,316 kW] Low Level Utility Heat Liquid Feed Heater 70,480
MBTU/Hr [45,526 kW] Demethanizer Reboiler 18 24,500 MBTU/Hr [15,826
kW] Totals 94,980 MBTU/Hr [61,352 kW] High Level Utility Heat
Demethanizer Reboiler 19 27,230 MBTU/Hr [17,589 kW] Specific Power
HP-Hr/Lb. Mole 1.795 [kW-Hr/kg mole] [2.950] *(Based on un-rounded
flow rates)
[0076] 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 significantly less power than the FIG. 3 embodiment
(improving the specific power by over 14%) and slightly less than
the FIG. 4 embodiment. However, the high level utility heat
required for the FIG. 5 embodiment of the present invention is
considerably lower than that of the FIG. 3 and FIG. 4 embodiments
(by about 13% and 17%, 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.
Other Embodiments
[0077] FIGS. 3 through 5 depict fractionation towers constructed in
a single vessel. FIGS. 6 through 8 depict fractionation towers
constructed in two vessels, absorber (rectifier) column 66 (a
contacting and separating device) and stripper (distillation)
column 20. In such cases, the overhead vapor (stream 43) from
stripper column 20 is split into two portions. One portion (stream
44) is routed to heat exchanger 52 to generate supplemental reflux
for absorber column 66. The remaining portion (stream 47) flows to
the lower section of absorber column 66 to be contacted by the cold
reflux (stream 82b) and the supplemental reflux (condensed liquid
stream 44b). Pump 67 is used to route the liquids (stream 46) from
the bottom of absorber column 66 to the top of stripper column 20
so that the two towers effectively function as one distillation
system. The decision whether to construct the fractionation tower
as a single vessel (such as demethanizer 20 in FIGS. 3 through 5)
or multiple vessels will depend on a number of factors such as
plant size, the distance to fabrication facilities, etc.
[0078] In accordance with this invention, it is generally
advantageous to design the absorbing (rectification) section of the
demethanizer to contain multiple theoretical separation stages.
However, the benefits of the present invention can be achieved with
as few as one theoretical stage, and it is believed that even the
equivalent of a fractional theoretical stage may allow achieving
these benefits. For instance, all or a part of the cold reflux
(stream 82b), all or a part of the condensed liquid (stream 44b),
and all or a part of streams 77a and 34a can be combined (such as
in the piping to the demethanizer) and if thoroughly intermingled,
the vapors and liquids will mix together and separate in accordance
with the relative volatilities of the various components of the
total combined streams. Such commingling of these streams shall be
considered for the purposes of this invention as constituting an
absorbing section.
[0079] In the examples shown, total condensation of streams 44a and
81b is illustrated in FIGS. 3 through 8. Some circumstances may
favor subcooling these streams, while other circumstances may favor
only partial condensation. Should partial condensation of either or
both 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.
[0080] 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 31a (FIGS. 3 and 6) or expanded cooled stream 31b (FIGS. 4,
5, 7, 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.
[0081] 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.
[0082] 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 52 and 53 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 distillation vapor
streams for process heat exchange, and the particular arrangement
of heat exchangers for heating the LNG streams and cooling the feed
gas stream, must be evaluated for each particular application, as
well as the choice of process streams for specific heat exchange
services.
[0083] 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. 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. 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.
[0084] 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 stream. 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.
[0085] The present invention provides improved recovery of C.sub.2
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.
[0086] In the examples given for the FIGS. 3 through 5 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 C.sub.3
components and heavier hydrocarbon components is desired.
[0087] 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.
* * * * *